Process for the preparation of cyclohexanedimethanol

ABSTRACT

This invention provides a process for preparing cyclohexanedimethanol comprising hydrogenating a cyclohexanedicarboxylic acid dialkyl ester by a fixed-bed continuous reaction in the presence of a preformed copper-containing catalyst under the conditions of reaction temperature of 200 to 280° C., hydrogen pressure of 185 to 300 kgf/cm 2  and hydrogen gas feed rate of 1 to 40 cm/s in terms of superficial linear velocity, the cyclohexanedicarboxylic acid dialkyl ester being prepared, typically, by ring hydrogenating of an aromatic dicarboxylic acid dialkyl ester in the presence of a preformed supported ruthenium catalyst by a fixed-bed continuous reaction.

FIELD OF THE INVENTION

The present invention relates to a process for preparingcyclohexanedimethanol (hereinafter referred to as “CHDM”). When used asthe diol component of polyester resins, polyurethane resins,polycarbonate resins or the like, CHDM is effective in improving theheat resistance, transparency, weatherability and molding properties ofthese resins. Particularly 1,4-cyclohexanedimethanol is drawingattention as a compound useful for improving the properties ofpolyethylene terephthalate.

BACKGROUND ART

Conventional processes for preparing CHDM generally comprise ringhydrogenation of aromatic dicarboxylic acid dialkyl ester to givecyclohexanedicarboxylic acid dialkyl ester (first reaction) andhydrogenation of its ester groups to give CHDM (second reaction).

Among these catalysts used for the respective reactions, known aseffective for the first reaction are palladium, nickel, ruthenium,rhodium and the like (U.S. Pat. No. 3,334,149, Japanese UnexaminedPatent Publications Nos. 163554/1979 and 192146/1994, U.S. Pat. No.5,286,898 and U.S. Pat. No. 5,399,742); and known as effective for thesecond reaction are copper-chromite, copper oxide/zinc oxide, copperoxide/titanium oxide, copper oxide/iron oxide and catalysts prepared bymodifying these copper-based catalysts with oxides of barium, magnesiumand zinc and reducing the modified catalysts for activation (U.S. Pat.No. 3,334,149, Japanese Unexamined Patent Publications Nos. 192146/1994and 196549/1995, U.S. Pat. No. 5,334,779 and U.S. Pat. No. 5,030,771,U.S. Pat. No. 4,929,777).

As the mode of reaction, a fixed-bed continuous reaction process isconsidered to be advantageous over a suspended catalyst reaction processin terms of productivity and yield. Herein, the fixed-bed continuousreaction process includes a reaction process (downflow-type process)wherein a preformed catalyst is placed into a pressure-resistantreactor, into which hydrogen and a raw material are supplied to the topof the reactor at a predetermined temperature and hydrogen pressure, andthe reaction product is withdrawn from the bottom of the reactor; and areaction process (upflow-type process) wherein hydrogen and a rawmaterial are supplied to the bottom of the reactor and the reactionproduct is removed from the top of the reactor. The suspended catalystreaction process comprises suspending a catalyst powder in an aromaticdicarboxylic acid diester or a cyclohexanedicarboxylic acid diester andsubjecting the suspension to a reaction with heating while beingpressurized with hydrogen.

As the examples of fixed-bed continuous reaction process, reported areprocesses comprising ring hydrogenation of a terephthalic acid dialkylester in the presence of a preformed supported ruthenium catalyst togive 1,4-cyclohexanedicarboxylic acid dialkyl ester in the firstreaction (Japanese Unexamined Patent Publications Nos. 163554/1979 and192146/1994).

Ruthenium catalysts are inexpensive compared with palladium catalysts,and exhibit a high activity even at a low pressure and a lowtemperature, but have the disadvantage of being likely to causeundesirable reactions involving high exothermic heat, such ashydrogenolysis of ester groups to hydroxymethyl group(s) or methylgroup(s), in addition to the hydrogenation of the aromatic ring.

It is advantageous to use a cyclohexanedicarboxylic acid diester as areaction solvent in order to avoid adverse effect due to reaction heat.However, the cyclohexanedicarboxylic acid diester used may be consumeddue to the above-mentioned side reactions, not only resulting in a lowyield but also, in extreme case, leading to a rapid generation of heatin a portion of the reactor which makes it difficult to continue thereaction. Therefore, Japanese Unexamined Patent Publication No.192146/1994 proposes the provision of a perforated plate in the reactorto improve the dispersibility of gas and liquid. However, even in thiscase, the concentration of terephthalic acid dialkyl ester in the feedto the reactor actually must be limited to an extremely low range of 5to 20 weight %, and a large amount of reaction product is subjected toreaction by recycling, resulting in a low yield based on theterephthalic acid dialkyl ester used and leading to a low productivity.

Examples of the fixed-bed continuous reaction process for the secondreaction are those disclosed in Japanese Unexamined Patent Publications.Nos. 196549/1995; 196560/1995; 196558/1995; 196559/1995; 188077/1995;188078/1995 and 188079/1995. These proposed processes are characterizedin that the reaction is conducted under gas phase conditions ofrelatively low hydrogen pressure. However, the processes entail variousdisadvantages such as loss of thermal energy in vaporizing the rawmaterial and necessity for removal of generated reaction heat in a gasphase of low heat conductivity, resulting in a need for complicatedequipments. Furthermore, high-boiling-point by-products are deposited onthe surface of catalyst, thereby markedly reducing the catalyst activityand thus necessitating frequent catalyst replacement or catalystregeneration treatment.

On the other hand, U.S. Pat. Nos. 3,334,149, 5,030,771 and 5,334,779 andJapanese Unexamined Patent Publication No. 192146/1994 disclose agas-liquid mixed phase reaction. However, the disclosed processesinclude various problems. For example, the feed rate (F/V) ofcyclohexanedicarboxylic acid dialkyl ester is as low as 1/h or less,leading to a low productivity per reactor. Alternatively,cyclohexanedicarboxylic acid dialkyl ester is fed as diluted with thereaction product, i.e. CHDM or the like to a concentration of about 16%by weight or less, consequently involving complicated equipment andcumbersome operation relating to the reactor and resulting in increasedby-products due to side reaction of CHDM.

DISCLOSURE OF THE INVENTION

An object of the present invention is to provide a process for preparinghigh-quality CHDM, which is capable of producing CHDM with safety in ahigh yield by a simplified procedure at a high productivity per reactorand capable of diminishing costs for equipment.

The present inventors found that in producing an alicyclic alcohol,particularly cyclohexanedimethanol, a high yield can be achieved usingsimplified equipment when cyclohexanedicarboxylic acid ester ishydrogenated under specific reaction conditions, and completed thepresent invention.

The present invention provides a process for preparingcyclohexanedimethanol represented by the formula (1)

the process comprising the step of hydrogenating cyclohexanedicarboxylicacid dialkyl ester represented by the formula (2)

wherein R represents an alkyl group having 1 to 12 carbon atoms or acycloalkyl group having 5 to 10 carbon atoms, and particularlyrepresents a straight- or branched-chain alkyl group having 1 to 4carbon atoms, by a fixed-bed continuous reaction in the presence of apreformed copper-containing catalyst under the condition of reactiontemperature of 200 to 280° C., hydrogen pressure of 185 to 300 kgf/cm²and hydrogen gas feed rate of 1 to 40 cm/s in terms of superficiallinear velocity.

In the specification and claims, superficial linear velocity has agenerally recognized meaning and is defined as a value obtained bydividing the hydrogen gas flow rate (cm³/s) by a cross-sectional area(cm²) of the tubular or columnar reactor in the case of a single-columnreactor, or as a value obtained by dividing the hydrogen gas flow rate(cm³/s) by a total cross-sectional area (cm²) of a plurality of tubes inthe case of a multi-tubular reactor.

The cyclohexanedicarboxylic acid dialkyl ester of the formula (2) canadvantageously be obtained by ring hydrogenation of benzenedicarboxylicacid dialkyl ester in the presence of a preformed supported rutheniumcatalyst.

In the present invention, accordingly, cyclohexanedimethanol canadvantageously be obtained by a process comprising the steps ofhydrogenating benzenedicarboxylic acid dialkyl ester (first reaction)and hydrogenating the obtained cyclohexanedicarboxylic acid dialkylester in the presence of a preformed copper-containing catalyst (secondreaction).

Thus, the present invention provides a process for preparingcyclohexanedimethanol represented by the formula (1)

the process comprising the steps of

(a) carrying out ring hydrogenation of an aromatic dicarboxylic aciddialkyl ester represented by the formula (3)

 wherein R represents an alkyl group having 1 to 12 carbon atoms or acycloalkyl group having 5 to 10 carbon atoms, and particularlyrepresents a straight- or branched-chain alkyl group having 1 to 4carbon atoms, in the presence of a preformed supported rutheniumcatalyst by a fixed-bed continuous reaction to givecyclohexanedicarboxylic acid dialkyl ester represented by the formula(2)

 wherein R is as defined above, and

(b) hydrogenating the cyclohexanedicarboxylic acid dialkyl esterobtained in step (a) and represented by the formula (2) by a fixed-bedcontinuous reaction in the presence of a preformed copper-containingcatalyst under the conditions of reaction temperature of 200 to 280° C.,hydrogen pressure of 185 to 300 kgf/cm² and hydrogen gas feed rate of 1to 40 cm/s in terms of superficial linear velocity.

According to one embodiment of this invention (embodiment I), CHDM canbe prepared by a process comprising the steps of hydrogenating the ringof a terephthalic acid dialkyl ester in the presence of a preformedsupported ruthenium catalyst by a fixed-bed continuous reaction to give1,4-cyclohexanedicarboxylic acid dialkyl ester (first reaction) andhydrogenating the obtained 1,4-cyclohexanedicarboxylic acid dialkylester by a fixed-bed continuous reaction in the presence of a preformedcopper-chromite catalyst (second reaction) wherein the hydrogenation ofthe second reaction is carried out under the conditions of reactiontemperature of 225 to 280° C., hydrogen pressure of 185 to 300 kgf/cm ²and hydrogen gas feed rate of 1 to 40 cm/s in terms of superficiallinear velocity.

This embodiment I has the feature of producing CHDM in a yieldcomparable to the prior art processes and furthermore extending the lifeof the catalyst.

According to the research of the present inventors, it has been revealedthat a fixed-bed continuous reaction in a downflow-type process with useof a multitubular pressure-resistant reactor can industrially producehigh-quality CHDM in high yields and at high productivity inconceivablefrom the prior art process.

Thus, according to another embodiment of the invention (embodiment II),there is provided a process for preparing CHDM, characterized in that itcomprises the steps of

hydrogenating an aromatic dicarboxylic acid dialkyl ester in thepresence of a preformed supported ruthenium catalyst by a fixed-bedcontinuous reaction to give a corresponding cyclohexanedicarboxylic aciddialkyl ester (first reaction) and

hydrogenating the obtained cyclohexanedicarboxylic acid dialkyl ester bya fixed-bed continuous reaction in the presence of a preformedcopper-containing catalyst under the conditions of reaction temperatureof 200 to 280° C., hydrogen pressure of 185 to 300 kgf/cm² and hydrogengas feed rate of 1 to 40 cm/s in terms of superficial linear velocity(second reaction),

wherein as the reactor for each step, a multitubular pressure-resistantreactor packed with each of the above specified catalysts is used, and

wherein hydrogen and a feed composed of each raw material and if desiredthe reaction product and/or a solvent are fed to the top of each reactorto carry out hydrogenation under a gas-liquid mixed phase condition, andexcess hydrogen and the reaction product are removed from the bottom ofeach reactor.

According to the research of the present inventors, it has been revealedthat hydrogenation of a cyclohexanecarboxylic acid mono-, di-, tri- ortetra-alkyl ester (particularly cyclohexanedicarboxylic acid dialkylester) in the presence of a preformed copper-containing catalyst insubstantially the same manner as in embodiment I above can produce acorresponding alicyclic alcohol (particularly CHDM), and that when thehydrogenation is carried out while feeding an aliphatic alcohol to saidreaction system, high-quality CHDM can be produced by an economicallyadvantageous process with use of a simplified equipment at highproductivity.

Thus, according to a further embodiment of the invention (embodimentIII), there is provided a process for preparing an alicyclic alcohol,particularly CHDM, represented by the formula (III-2)

wherein n is as defined in the formula (III-1), the process beingcharacterized in that it comprises the step of hydrogenatingcyclohexanecarboxylic acid ester represented by the formula (III-1)

wherein R¹ is a straight- or branched-chain alkyl group having 1 to 4carbon atoms, and n is an integer of 1 to 4, particularly 2, with theproviso that the R¹ groups which are n in number may be the same ordifferent, by a fixed-bed continuous reaction in the presence of apreformed copper-containing catalyst under the conditions of reactiontemperature of 200 to 280° C., hydrogen pressure of 185 to 300 kgf/cm²and hydrogen gas feed rate of 1 to 40 cm/s in terms of superficiallinear velocity,

wherein hydrogen, the cyclohexanecarboxylic acid ester and an aliphaticalcohol having 1 to 4 carbon atoms are fed to a reactor packed with saidcatalyst.

This embodiment III has the feature of improving the yield of CHDM bythe use of said aliphatic alcohol even without the use of a multitubularpressure-resistant reactor described with respect to embodiment II.

According to a still further embodiment of the invention (embodimentIV), it is preferable that the cyclohexanedicarboxylic acid dialkylester, which is an intermediate for preparing CHDM, is prepared by aprocess using a ruthenium catalyst having specific physical properties.Said process is a process for preparing cyclohexanedicarboxylic aciddialkyl ester represented by the formula (IV-2)

wherein R¹ and R² are the same or different and each represents an alkylgroup having 1 to 4 carbon atoms by hydrogenating an aromaticdicarboxylic acid dialkyl ester represented by the formula (IV-1)

wherein R¹ and R² are as defined above in the presence of a preformedsupported ruthenium catalyst by a fixed-bed continuous reaction,characterized in that the chlorine content of the preformed supportedruthenium catalyst is 500 ppm or less.

BRIEF DESCRIPTION DRAWING

FIG. 1 a schematic view showing the reactor used in the examplesillustrating embodiment II.

DETAILED DESCRIPTION OF THE INVENTION

The above embodiments I to IV will be described below in detail.

Embodiment I

The present inventors conducted extensive research to overcome theforegoing prior art problems and found the following. In carrying outthe ring hydrogenation of terephthalic acid dialkyl ester and thehydrogenation of the ester groups of 1,4-cyclohexanedicarboxylic aciddialkyl ester by a fixed-bed continuous reactor, the selection of aspecific temperature range, a hydrogen pressure range and hydrogen gasfeed rate, especially in the hydrogenation of the ester groups,remarkably suppresses the formation of high-boiling point by-products,consequently leading to a significantly extended life of thecopper-chromite catalyst. The present invention was completed based onthis finding.

Thus, the process for preparing CHDM according to embodiment I ischaracterized in that it comprises the steps of hydrogenating the ringof a terephthalic acid dialkyl ester in the presence of a preformedsupported ruthenium catalyst by a fixed-bed continuous reaction to give1,4-cyclohexanedicarboxylic acid dialkyl ester (first reaction) andhydrogenating the obtained 1,4-cyclohexanedicarboxylic acid dialkylester in the presence of a preformed copper-chromite catalyst by afixed-bed continuous reaction to give CHDM (second reaction), whereinthe hydrogenation in the second reaction is carried out under theconditions of a reaction temperature of 225 to 280° C., hydrogenpressure of 185 to 300 kgf/cm² and hydrogen feed rate of 1 to 40 cm/s interms of superficial linear velocity.

The reactor useful for the fixed-bed continuous reaction of the presentinvention, for both of the first reaction and the second reaction, maybe of a single column type, or of a multitubular type which comprises aplurality of tubes having a small interior diameter and arranged inparallel (shell-and-tube reactor).

First Reaction

The terephthalic acid dialkyl ester for use as the starting material isa diester prepared by esterifying terephthalic acid, as an acidcomponent, with a straight- or branched-chain aliphatic alcohol having 1to 12 carbon atoms or an alicyclic alcohol having 5 to 10 carbon atoms,particularly a straight- or branched-chain aliphatic alcohol having 1 to4 carbon atoms, as an alcohol component, in the conventional manner.

While any of primary, secondary and tertiary alcohols can be used as thealcohol component, primary and secondary alcohols are preferred.Specific examples are methanol, ethanol, n-propanol, iso-propanol,n-butanol, iso-butanol, n-hexanol, cyclohexanol, n-octanol,2-ethylhexanol, n-decanol and lauryl alcohol.

Typical examples of terephthalic acid dialkyl esters are dimethylterephthalate, diethyl terephthalate, di-n-propyl terephthalate,di-n-butyl terephthalate and di-2-ethylhexyl terephthalate. Among them,preferred is dimethyl terephthalate which is prepared using methanol asthe alcohol component and commercially available.

The preformed supported ruthenium catalyst useful in the invention canbe any moldings, such as tablets, pellets, cylinders and spheres, ofconventional supported ruthenium catalysts which are known as catalystsfor hydrogenating aromatic rings. Examples of useful support arealumina, silica, titania, magnesia and zirconia, among which alumina ispreferred.

The amount of ruthenium to be deposited on a support is recommendably0.05 to 10% by weight, and preferably 0.1 to 5% by weight, based on thesupport. If the amount is less than 0.05% by weight, a pronouncedly lowactivity is exhibited and hence it is impractical. If the amount is morethan 10% by weight, only the cost will increase without noticeablyimproving the catalyst activity and furthermore marked separation of thedeposited ruthenium from the support takes place, hence impractical.

These preformed catalysts can be used as such, or can be used in thereaction after effecting a suitable activation treatment, such asreduction, in the conventional manner.

The shape of the preformed supported ruthenium catalyst is notspecifically limited, but generally the catalysts of cylindrical shapewhich are commercially readily available are used. Their size can bedetermined according to the interior diameter of the reactor. Usuallypreferred are cylindrical catalysts having a diameter of 2 to 6 mm and aheight of 2 to 6 mm.

In this embodiment I, it is preferable to use the specific rutheniumcatalyst to be described later in embodiment IV.

Generally, the higher the hydrogen partial pressure, the more smoothlythe first reaction proceeds. If the hydrogen pressure becomes higherthan necessary, a special pressure-resistant reactor is required, andits use is uneconomical. Practically, it is preferable that the hydrogenpressure is in the range of 5 to 100 kgf/cm², particularly 30 to 100kgf/cm².

Basically, the pressure in the reaction system is a sum of said hydrogenpressure and vapor pressures of the starting material and the productand partial pressures of methane gas and the like that are formed asby-products. However, the vapor pressures of the starting material andthe product and partial pressures of the gas are almost negligible andtherefore the reaction pressure is substantially equal to the hydrogenpressure.

The reaction temperature is, for example, in the range of 80 to 200° C.,recommendably 90 to 160° C. At a reaction temperature of below 80° C.,the reaction rate becomes markedly slow, whereas at a reactiontemperature of above 200° C., side reactions take place preferentially,and hence impractical.

The first reaction can be conducted without use of a solvent because thestarting material terephthalic acid dialkyl ester is fed as such when itis liquid, or fed as melted when it is a solid. However, a solvent ispreferably used when terephthalic acid dialkyl ester used as the rawmaterial is cumbersome to handle owing to the high melting point thereofor when it is necessary to facilitate the removal of reaction heat.

When the reaction solvent is used, the kind of the solvent is notspecifically limited insofar as it does not adversely affect thereaction. Specific examples thereof include a1,4-cyclohexanedicarboxylic acid dialkyl ester and an alcohol whichcorrespond to the terephthalic acid dialkyl ester used as the rawmaterial in the reaction. Particularly, the most preferred solvent is1,4-cyclohexanedicarboxylic acid dialkyl ester.

Examples of the 1,4-cyclohexanedicarboxylic acid dialkyl ester are thereaction products obtained in the first reaction, such as dimethyl1,4-cyclohexanedicarboxylate, diethyl 1,4-cyclohexanedicarboxylate, anddipropyl 1,4-cyclohexanedicarboxylate.

Examples of the alcohol are, for example, those corresponding toterephthalic acid dialkyl ester used as the raw material, such asmethanol, ethanol, n-propanol, iso-propanol, n-butanol, iso-butanol,n-hexanol, cyclohexanol, n-octanol, 2-ethylhexanol, etc.

When the reaction solvent is used, the amount of the solvent to be usedis suitably selected and is adjusted in such a manner that theconcentration of terephthalic acid dialkyl ester in the system is 5 to80% by weight, preferably 10 to 50% by weight. Below 5% by weight, a lowproductivity results, whereas over 80% by weight, there would be noadvantage of using the solvent.

The mode of the fixed-bed hydrogenation process in the first reactionmay be either of a downflow type or of an upflow type wherein the rawmaterial and hydrogen are supplied to the top or to the bottom of thereactor packed with the above-mentioned catalyst. However, the mode ofthe downflow type is usually preferable from the standpoint of thecatalyst life.

The feed rate of terephthalic acid dialkyl ester to be used as the rawmaterial is preferably 0.1 to 5/h, more preferably 0.2 to 3/h, in termsof F/V (wherein F represents the feed rate (liter/h) of terephthalicacid dialkyl ester and V is the volume (liter) of the catalyst bed inthe reactor.

The feed rate of hydrogen is recommendably 1 to 40 cm/s, preferably 2 to10 cm/s, in terms of superficial linear velocity under the reactionconditions.

In this first reaction, the terephthalic acid dialkyl ester used as theraw material is substantially quantitatively hydrogenated, so that withrespect to the reaction mixture discharged from the fixed-bed reactor,usually more than 95 weight % of the terephthalic acid dialkyl esterused has been converted to the desired 1,4-cyclohexanedicarboxylic aciddialkyl ester.

The thus obtained reaction mixture as such or the1,4-cyclohexanedicarboxylic acid dialkyl ester isolated from thereaction mixture is fed, as the raw material in the second reaction, toa reactor for the second reaction. From the viewpoint of operation, itis advantageous that the reaction mixture obtained in the first reactionas such is used as the raw material of the second reaction.

Second Reaction

This embodiment I is mainly characterized in that the above reactionproduct of the first reaction is hydrogenated in the second reactionunder the conditions of the specific reaction temperature (225-280° C.),the specific hydrogen pressure (185 to 300 kgf/cm²) and the specifichydrogen gas superficial linear velocity (1 to 40 cm/s), whereby thelife of the copper-chromite catalyst used is extended and the desiredCHDM can be produced in a high yield.

The catalyst to be used for hydrogenation of the ester groups in thesecond reaction is a preformed copper-chromite catalyst. The catalystmay contain, as a promoter, oxides of barium or manganese to enhance thecatalyst activity and/or to prevent the sintering of the catalyst.Further usable are those molded after addition of various binders inorder to impart an improved strength to the catalyst.

Specific examples of such catalysts include, for example, commerciallyavailable preformed copper-chromite catalysts of the so-called Adkinstype. Preferred are preformed multi-element type copper-chromitecatalysts containing one or more promoters such as barium oxide ormanganese oxide.

Generally, the copper-chromite catalyst contains copper in an amount,calculated as CuO, of 20-80 wt. %, preferably 30-70 wt. % and chromiumin an amount, calculated as Cr₂O₃, of 15-70 wt. %, preferably 40-60 wt.%. The above Adkins type catalyst preferably contains copper in anamount, calculated as CuO, of 30-60 wt. % and chromium in an amount,calculated as Cr₂O₃, of 30-60 wt. %. These copper-containing catalystspreferably contain the above promoter(s) in an amount of up to 10 wt. %,calculated as the metal oxide. In this embodiment I, the content ofbarium or manganese is, for example, 0.5 to 10% by weight, calculated asbarium oxide or manganese oxide.

In order to control exothermic heat abruptly generated at the start ofthe reaction and effectively exhibit a catalyst activity, it iseffective that the copper-chromite catalyst is subjected to preliminaryreduction treatment by conventional methods.

The shape of copper-chromite catalysts is not specifically limited, butcatalysts of cylindrical shape which are commercially readily availableare usually recommended. The size thereof can be determined according tothe interior diameter of the reactor to be used, and usually preferredare cylindrical catalysts having a diameter of 2 to 6 mm and a height of2 to 6 mm.

The hydrogenation conditions in the second reaction can be suitablyselected depending on the kind of 1,4-cyclohexanedicarboxylic aciddialkyl ester to be used as the raw material, and are generally asfollows.

The reaction temperature is in the range of 225 to 280° C., preferably240 to 265° C. At a reaction temperature of lower than 225° C., amarkedly low reaction rate is exhibited, and a high-boiling-point estercompounds are produced as by-products in a large amount, whereas at areaction temperature of higher than 280° C., decomposition reaction andcondensation reaction would markedly occur, and in either case thecatalyst has a short life so that such temperature conditions areimpractical. Incidentally, when a specific reactor is used as inembodiment II to be described below or when an alcohol is used as inembodiment III to be described below, reaction temperature of about 200°C. may be used.

The higher the hydrogen pressure, the more smoothly the second reactionproceeds. Practically, however, it is preferable to select the hydrogenpressure in the range of 185 to 300 kgf/cm² particularly 200 to 250kgf/cm². At a hydrogen pressure of below 185 kgf/cm², it is difficult toachieve a suitable reaction rate and high-boiling-point ester compoundsare produced as by-products in a large amount so that the catalyst lifebecomes short, whereas at a hydrogen pressure of above 300 kgf/cm²,special pressure-resistant reactor is required, and thereforeuneconomical. The pressure of the whole reaction system is basically thesame as, or slightly higher than, the hydrogen pressure.

A reaction solvent is usually unnecessary, because1,4-cyclohexanedicarboxylic acid dialkyl ester used as the raw materialis usually liquid. However, a suitable solvent may be used when1,4-cyclohexanedicarboxylic acid dialkyl ester is difficult to handleowing to the high melting point thereof or when facilitated removal ofreaction heat is required.

The mode of the fixed-bed hydrogenation process in the second reactionmay be either of a downflow type or of an upflow type wherein the rawmaterial and hydrogen are supplied to the top or the bottom of thereactor packed with the above preformed catalysts. However, the mode ofthe downflow type is usually preferable from the standpoint of thecatalyst life.

The feed rate of 1,4-cyclohexanedicarboxylic acid dialkyl ester isrecommendably 0.1 to 5/h, preferably 0.2 to 2/h, in terms of F/V(wherein F represents the feed rate (liter/h) of1,4-cyclohexanedicarboxylic acid dialkyl ester and V is the volume(liter) of the catalyst bed in the reactor.

The feed rate of hydrogen is recommendably 1 to 40 cm/s, preferably 2 to20 cm/s, in terms of superficial linear velocity under the reactionconditions. At a superficial linear velocity of lower than 1 cm/s, thereaction rate decreases and the quantity of by-products increases. Onthe other hand, at a feed rate of higher than 40 cm/s, there is nofurther appreciable improvement in the reaction rate, resulting ineconomical disadvantage, and the duration of activity and catalyststrength would decrease.

In this reaction, the ester groups of the 1,4-cyclohexanedicarboxylicacid dialkyl ester used as the raw material are substantiallyquantitatively converted to hydroxymethyl group by hydrogenation, sothat the reaction mixture discharged from the fixed-bed reactor usuallycontains the desired 1,4-cyclohexanedimethanol in an amount of more than95% by weight based on the weight of the reaction mixture excluding theformed alcohol and also excluding the solvent(s) when the solvent(s)is(are) used.

The thus-obtained CHDM can be purified by conventional methods such asdistillation.

Embodiment II

According to embodiment II of the present invention, there is provided aprocess characterized in that the process comprises the steps of:

(a) hydrogenating an aromatic dicarboxylic acid dialkyl ester with useof a preformed supported ruthenium catalyst by a fixed-bed continuousreaction to give a corresponding cyclohexanedicarboxylic acid dialkylester (first reaction), and

(b) hydrogenating the obtained cyclohaxanedicarboxylic acid dialkylester by a fixed-bed continuous reaction in the presence of a preformedcopper-containing catalyst under the conditions of reaction temperatureof 200 to 280° C., hydrogen pressure of 185 to 300 kgf/cm² and hydrogengas feed rate of 1 to 40 cm/s in terms of superficial linear velocity toprepare cyclohexanedimethanol (second reaction),

wherein a multitubular pressure-resistant reactor packed with each ofthe above specified catalysts is used as the reactor in each of thesteps (a) and (b), and wherein hydrogen and each raw material are fed tothe top of each reactor to effect hydrogenation under a gas-liquid mixedphase condition, and excess hydrogen and the reaction product arewithdrawn from the bottom of each reactor.

In other words, according to said embodiment II of the invention, thereis provided a process for preparing cyclohexanedimethanol, the processcomprising the steps of:

(a) continuously feeding an aromatic dicarboxylic acid dialkyl ester andhydrogen to the top of a multitubular pressure-resistant reactor packedwith a preformed supported ruthenium catalyst to effect hydrogenationunder a gas-liquid mixed phase condition, and removing excess hydrogenand the corresponding cyclohexanedicarboxylic acid dialkyl ester fromthe bottom of said reactor (first reaction), and

(b) continuously feeding the cyclohexanedicarboxylic acid dialkyl esterobtained in step (a) above and hydrogen to the top of a multitubularpressure-resistant reactor packed with a preformed copper-containingcatalyst to effect hydrogenation under a gas-liquid mixed phasecondition and under the conditions of reaction temperature of 200 to280° C., hydrogen pressure of 185 to 300 kgf/cm² and hydrogen gas feedrate of 1 to 40 cm/s in terms of superficial linear velocity, andremoving excess hydrogen and the resulting cyclohexanedimethanol fromthe bottom of said reactor (second reaction).

Desirably the following conditions are employed in embodiment II.

In the first reaction, namely in step (a), it is preferable that thehydrogen pressure is 30 to 100 kgf/cm², the reaction temperature is 120to 180° C., the superficial linear velocity of hydrogen gas is 1 to 15cm/s, particularly 1 to 10 cm/s, and the concentration of the aromaticdicarboxylic acid dialkyl ester in the feed to the reactor is at least30% by weight.

In the second reaction, namely step (b), it is preferable that thehydrogen pressure is 185 to 300 kgf/cm², the reaction temperature is 200to 280° C., the superficial linear velocity of hydrogen gas is 1 to 40cm/s, more preferably 5 to 30 cm/s, and the concentration ofcyclohexandedicarboxylic acid dialkyl ester in the feed to the reactoris at least 90% by weight.

In the second reaction, cyclohexanedicarboxylic acid dialkyl ester ispreferably fed at a feed rate (F/V) of 1.1 to 3.0/h (F/V=feed rate perhour relative to the volume of the catalyst bed in the reactor; F is afeed rate (liter/h) of the cyclohexanedicarboxylic acid dialkyl esterand V is a volume (liter) of the catalyst bed in the reactor).

Further, the reactor to be used in embodiment II preferably comprises ashell and a plurality of tubes arranged in parallel and housed in theshell, and a heat transfer medium is passed through the shell to heat orcool the tubes. Especially, it is preferable that the shell for thereactor for the first reaction is divided into at least two zones, andthat the temperature of each of the zones can be independentlycontrolled, whereby the temperature difference within the reactor willbe not greater than 50 degrees (° C.), particularly not greater than 30degrees (° C.).

Preferably the reactor for step (a), namely the first reaction is heatedor cooled with a shell having at least two zones, and the heat transfermedia flowing through the zones are independently heated or cooled sothat the temperature difference in the reactor will be not greater than50 degrees (° C.).

The reactors to be used for the fixed-bed continuous reaction accordingto embodiment II, for both of the first and second reactions, aremultitubular reactors, in which a plurality of tubes having a smallinterior diameter are arranged in parallel, and which are equipped withmeans for heating and cooling a heat transfer medium flowing through ashell thereof.

The interior diameter of each tube constituting the multitubular reactoris preferably 2.5 to 10 cm, in particular 3 to 6 cm. If tubes eachhaving an interior diameter of less than 2.5 cm are used, the number oftubes required for achieving the desired production output will beexcessive and additionally the productivity will be lowered. Conversely,if tubes each having an interior diameter of larger than 10 cm are used,the gas-liquid dispersion efficiency is lowered, making it difficult toachieve the high productivity, high quality and high yield ascontemplated in the present invention.

Preferably, each tube has a length of 3 to 15 m, particularly 5 to 10 m.The use of tubes less than 3 m or over 15 m in length results inpronouncedly reduced productivity and lowered yield of the desired CHDM.

The number of tubes is advantageously at least 10, particularly 10 to2000, in view of the cost for manufacturing the reactor. However, itshould be noted that the upper limit of the number of tubes is notparticularly restricted but may be suitably selected depending on thedesired production output.

The shell through which the heat transfer medium passes may be of anon-partition type or of a multi-partition type internally separatedinto plural zones, the temperature each of which can be independentlycontrolled. Especially, the reactor for the first reaction preferablyhas a shell of a multi-partition type with at least 2 zones, preferably3 to 6 zones. If the first reaction is conducted under a temperaturecontrol in a single-zone cell, there is a tendency that an intensiveexothermic reaction can not be controlled, causing various sidereactions and resulting in a low-quality product and a diminished yield.

As the mode of the fixed-bed continuous reaction according to embodimentII, a gas-liquid mixed phase downflow-type process is carried out,wherein hydrogen gas and raw material are fed to the top of the reactorpacked with the preformed catalyst, and the reaction product and theexcess hydrogen are removed from the bottom of the reactor, and whereinthe reaction is carried out under the reaction temperature and hydrogenfeed rate conditions such that the reaction temperature is lower thanthe dew point of at least one of the raw material and the product.

Problems are raised if the reaction of embodiment II is conducted by anupflow-type process wherein a raw material and hydrogen are fed from thebottom of the reactor or by a countercurrent-type process wherein, forexample, raw material is fed to the top of the reactor and hydrogen issupplied to the bottom of the reactor. In such case, the strength ofcatalyst is adversely affected by the friction caused due to themovement of the catalysts, necessitating frequent replacement ofcatalysts. Especially, ruthenium metal is separated from the catalystused in the first reaction and is lost, whereby the activity of thecatalyst is reduced within a short period of time.

Also, under the gas phase conditions in which the reaction temperatureis above a dew points of the raw material and the reaction product,high-boiling-point by-products are deposited on the surface of catalyst,significantly reducing the catalyst activity, thereby necessitatingfrequent replacement of catalysts and regeneration treatment thereof.Furthermore, under such conditions, it is necessary to remove thereaction heat in a gas phase of low heat conductivity so that extremelycomplicated equipment is required.

First Reaction

In the first reaction according to embodiment II, an aromaticdicarboxylic acid dialkyl ester is hydrogenated in the presence of apreformed supported ruthenium catalyst to give the correspondingcyclohexanedicarboxylic dialkyl ester.

The aromatic dicarboxylic acid dialkyl ester to be used as the rawmaterial is a diester prepared by esterifying terephthalic acid,isophthalic acid or phthalic acid with a monohydric aliphatic alcoholhaving 1 to 4 carbon atoms in the conventional manner. Among thesediesters, preferred is a compound represented by the formula (II-1)

wherein R is an alkyl group having 1 to 4 carbon atoms.

Specific examples of the above alcohol ROH are methanol, ethanol,n-propanol, iso-propanol, n-butanol, iso-butanol, etc.

Thus, examples of the aromatic dicarboxylic acid dialkyl ester aredimethyl terephthalate, diethyl terephthalate, di-iso-propylterephthalate, di-n-butyl terephthalate, dimethyl isophthalate, diethylisophthalate, di-iso-propyl isophthalate, di-n-butyl isophthalate,dimethyl phthalate, diethyl phthalate, di-iso-propyl phthalate,di-n-butyl phthalate, etc. Among them, the most preferred arecommercially available dimethyl terephthalate and dimethyl isophthalateprepared from methanol.

Useful preformed supported ruthenium catalyst include conventionalpreformed supported ruthenium catalysts which are known as catalysts forhydrogenating aromatic rings.

Useful carriers or support can be any of alumina, silica, titania,magnesia, zirconia, silicon carbide and the like, among which alumina ispreferred.

The amount of ruthenium to be deposited on the support is preferably0.05 to 10% by weight, more preferably 0.1 to 5% by weight, based on thesupport.

As to the shape of the preformed supported ruthenium catalyst, generallyrecommended are those of cylindrical shape which are commerciallyreadily available.

The size of the preformed catalyst can be determined according to theinterior diameter of the reactor to be used. Usually preferred arecylindrical catalysts having a diameter of 2 to 6 mm and a height of 2to 6 mm.

These preformed catalysts can be used as such or can be subjected to thereaction after being suitably activated as by reduction.

The specific ruthenium catalyst to be described later in embodiment IVmay be used in this embodiment II as well.

Generally, the higher the hydrogen partial pressure, the more smoothlythe hydrogenation reaction proceeds. If the hydrogen pressure is higherthan necessary, a special pressure-resistant reactor is required, andhence uneconomical.

The pressure to be employed in the first reaction is preferably in therange of 30 to 100 kgf/cm². At a pressure of below 30 kgf/cm², a lowreaction rate results, hence undesirable. Above 100 kgf/cm², in additionto the above-mentioned equipment problem, various disadvantages result.For example, side reactions which involve large amount of exothermicheat tend to occur, such as hydrogenolysis of ester groups tohydroxymethyl group(s) or to methyl group(s). Further, low yields anddifficulty in the control of reaction are entailed.

A preferred reaction temperature is in the range of 120 to 180° C. At areaction temperature of below 120° C., a reaction rate is markedly low,whereas at a reaction temperature of above 180° C., side reactionspreferentially occur. Hence the reaction temperature outside said rangeis not practical.

The first reaction is preferably carried out utilizing theabove-mentioned mode of heating or cooling the heat transfer medium tomaintain a temperature difference within 50 degrees (° C.), preferablywithin 30 degrees (° C.), in the longitudinal direction of tube withinthe reactor (namely in the direction of fluid flow). If there is atemperature difference beyond this range in the tubes of the reactor,there take place side reactions involving high exothermic heat, such ashydrogenolysis of ester groups to hydroxymethyl group(s) or to methylgroup(s), which tend to induce lowered yield and markedly decreasedproductivity.

The aromatic dicarboxylic acid dialkyl ester can be continuously fed tothe top of the reactor either singly or in the form of a mixture withthe reaction product obtained in the first reaction and mainlycontaining cyclohexane-dicarboxylic acid dialkyl ester. In the lattercase, high-melting-point aromatic dicarboxylic acid dialkyl esterbecomes easily melted and the reaction heat is easily controlled.

As to the proportions of the aromatic dicarboxylic acid dialkyl esterand the first reaction's reaction product, since the above-mentionedmultitubular reactor is used, the higher the concentration of thearomatic dicarboxylic acid dialkyl ester is, the more preferable fromthe standpoint of productivity and suppression of the formation ofby-products. Accordingly, it is recommendable that the concentration ofthe aromatic dicarboxylic acid dialkyl ester in the mixture is at least30% by weight, more preferably at least 40% by weight.

At a low concentration of less than 30% by weight, a low productivityresults, and the side reaction of cyclohexanedicarboxylic acid dialkylester circulating through the reaction system takes place, causing thedecrease in yield and selectivity. The aromatic dicarboxylic aciddialkyl ester can be used as diluted with a solvent other than theproduct of the first reaction, which solvent will not adversely affectthe reaction. The use of such solvent, however, requires additionalprocedures for separating and recovering the solvent and does not give aparticularly favorable result.

Recommended feed rate of the aromatic dicarboxylic acid dialkyl ester is0.1 to 5/h, preferably 0.2 to 3/h, and particularly preferable feed rateis 0.5 to 3/h, in terms of F/V (=feed rate of the aromatic dicarboxylicacid dialkyl ester per hour relative to the volume of the catalyst bedin the reactor). At a feed rate of lower than this range, productivityis lower, whereas at a feed rate of higher than this range, theresulting reaction product contains a large amount of unreacted aromaticdicarboxylic acid ester and is not suitable as the raw material in thesecond reaction. Herein, F stands for the feed rate (liter/h) of thearomatic dicarboxylic acid dialkyl ester per hour, and V represents thevolume (liter) of catalyst bed in the reactor.

The feed rate of hydrogen is 1 to 15 cm/s, preferably 1 to 10 cm/s interms of superficial linear velocity under the reaction condition. At asuperficial linear velocity lower than this range, it is difficult toachieve an effective contact between the gas and liquid on the surfaceof the catalyst with the result that the reaction rate is decreased andthe quantity of by-products due to the above-mentioned side reactionsincreases. On the other hand, at a feed rate higher than the aboverange, there is no further appreciable improvement in the reaction,resulting in economical disadvantage. Hence it is undesirable.

Second Reaction

The second reaction according to embodiment II is a reaction tohydrogenate the hydrogenation product of the aromatic dicarboxylic aciddialkyl ester of the formula (II-1) obtained in the first reaction,namely cyclohexanedicarboxylic acid dialkyl ester, using a preformedcopper-containing catalyst to give a corresponding CHDM.

Examples of useful preformed copper-containing catalysts includeconventional preformed copper-containing catalysts having ability toreduce esters, such as copper-chromite, copper oxide/zinc oxide, copperoxide/iron oxide, copper oxide/aluminum oxide, and these catalysts whichcontain oxide(s) of barium, manganese, aluminum, zinc or magnesium as apromoter.

Further usable are those molded after addition of various binders inorder to maintain an improved strength of the catalyst, and supportedcatalysts prepared by depositing said oxides on a support such asalumina, silica, silica-alumina, silicon carbide, zirconia, titania orzinc oxide.

Among the above copper-containing catalysts, particularly preferred arecommercially available copper-chromite catalysts of the so-called Adkinstype, copper oxide/zinc oxide catalysts and these catalysts whichcontain one or more promoters such as barium oxide or manganese oxide.

Generally, the above copper-containing catalysts contain copper in anamount, calculated as CuO, of 20-80 wt. %, preferably 30-70 wt. % andchromium in an amount, calculated as Cr₂O₃, of 15-70 wt. %, preferably40-60 wt. %. The above Adkins type catalyst preferably contains copperin an amount, calculated as CuO, of 30-60 wt. % and chromium in anamount, calculated as Cr₂O₃, of 30-60 wt. %. Other copper/metal oxidecatalysts mentioned above preferably contains copper in an amount,calculated as CuO, of 20-95 wt. % and other metal in an amount,calculated as the metal oxide, of 5-80 wt. %. These copper-containingcatalysts preferably contain said one or more promoters mentioned abovein an amount of up to 10 wt. %, calculated as said metal oxide. In thecase of the supported catalysts, the percentage values are thosecalculated after excluding the amount of the support.

In order to control an exothermic heat abruptly generated at the startof reaction and effectively exhibit a catalyst activity, it is effectivethat the preformed copper-containing catalyst is subjected topreliminary reduction treatment.

The preliminary reduction treatment can be conducted in a conventionalmanner under a stream of hydrogen-nitrogen gas mixture at an atmosphericpressure or elevated pressure at a temperature in the range of 150 to300° C. while gradually increasing the concentration of hydrogen.

As to the shape of copper-containing catalysts, catalysts of cylindricalshape which are commercially readily available are usually recommended.

The size thereof can be determined according to the interior diameter ofthe reactor to be used, and usually preferred are cylindrical catalystshaving a diameter of 2 to 6 mm and a height of 2 to 6 mm.

The hydrogen pressure is preferably 185 to 300 kgf/cm², more preferably200 to 250 kgf/cm². If the hydrogen pressure is below this range,various disadvantages are entailed, in addition to reduced reaction rateand decreased productivity, including increase in the quantity ofhigh-boiling-point by-products such as ether compounds and wax estersand lowered yields and selectivity. A lower hydrogen pressure is alsoundesirable from the viewpoint of the duration of catalytic acitivityand strength of the catalyst. Conversely, the use of a hydrogen pressureof greater than the above range only increases the equipment costwithout further appreciable improvement in the reaction rate andselectivity, hence undesirable.

The reaction temperature is preferably 200 to 280° C., more preferably225 to 280° C. At a reaction temperature of below this temperaturerange, a markedly low reaction rate is exhibited, whereas at a reactiontemperature higher than the above range, side reactions predominantlyoccur. Hence the reaction temperature outside said range tends to beimpractical.

In the second reaction, usually the reaction product obtained in thefirst reaction is continuously supplied, as it is, to the top of thereactor. The amount of unreacted aromatic dicarboxylic acid dialkylester in the reaction product obtained by the first reaction ispreferably controlled to 5% by weight or less for producing the desiredCHDM in a high yield.

Also a portion of the reaction product obtained in the second reactionand predominantly containing CHDM may be fed as mixed with the reactionproduct obtained in the first reaction, but the concentration ofcyclohexanedicarboxylic acid dialkyl ester in such feed to the secondreaction reactor is preferably at least 90% by weight in view of theproductivity and the increase in the amount of by-products due torelatively prolonged exposure of the desired CHDM to the reactionenvironment.

The feed rate of the reaction product obtained in the first reaction andto be fed to the reactor for the second reaction is preferably 1.1 to3.0/h, in terms of F/V based on cyclohexanedicarboxylic acid dialkylester. If the F/V value is less than 1.1/h, productivity is low witheconomical disadvantage, and also the yield of the desired CHDM isreduced due to side reactions, hence undesirable. Conversely, at a feedrate in excess of 3.0/h, the obtained reaction product containsunreacted cyclohexanedicarboxylic acid dialkyl ester which brings aboutpronounced decrease in the yield and purity.

Recommended feed rate of hydrogen gas is 1 to 40 cm/s, preferably 5 to30 cm/s, particularly 10 to 20 cm/s, in terms of superficial linearvelocity. The hydrogen gas superficial linear velocity has an extremelyremarkable effect on the second reaction, and at a velocity of less than5 cm/s, particularly less than 1 cm/s, the reaction rate significantlyreduces, making it difficult to achieve the high productivity ascontemplated in the present invention. On the other hand, at asuperficial linear velocity of more than 30 cm/s, particularly more than40 cm/s, a further advantage in terms of the reaction is not achieved,and the duration of activity and strength of the catalyst decreases, andonly the costs for equipment such as a hydrogen circulator and forenergy increases, hence undesirable.

Embodiment III

According to the present inventors' research, it has been found thatwhen cyclohexane mono-, di-, tri- or tetra-carboxylic acid lower alkylester is used as the raw material, together with an aliphatic alcoholhaving 1 to 4 carbon atoms, in the same manner as in the second reactionaccording to embodiment I, the corresponding alicyclic alcohol, i.e.cyclohexane-mono-, di-, tri or tetramethanol can be produced.

Thus, according to embodiment III of the present invention, there isprovided a process for preparing an alicyclic alcohol, particularlycyclohexanedimethanol (CHDM), represented by the formula (III-2)

wherein n is as defined in the formula (III-I), the process comprisingthe steps of hydrogenating a cyclohexanecarboxylic acid esterrepresented by the formula (III-1)

wherein R¹ is a straight- or branched-chain alkyl group having 1 to 4carbon atoms, and n is an integer of 1 to 4, particularly 2 providedthat the R¹ groups which are n in number may be the same or different,in the presence of a preformed copper-containing catalyst by a fixed-bedcontinuous reaction, wherein hydrogen, the cyclohexanecarboxylic acidester of the formula (III-1) and aliphatic alcohol having 1 to 4 carbonatoms are fed to a reactor packed with said catalyst.

In embodiment III, preferably hydrogen, a cyclohexanecarboxylic acidester of the formula (III-1) and an aliphatic alcohol having 1 to 4carbon atoms are fed to the top of the reactor packed with saidpreformed copper-containing catalyst, and excess hydrogen, the reactionproduct of the formula (III-2) and the aliphatic alcohol having 1 to 4carbon atoms are removed from the bottom of the reactor.

A preferred amount of the aliphatic alcohol having 1 to 4 carbon atomsto be fed to the reaction system is 1 to 100% by weight based on theester used as the raw material.

Desirably the cyclohexanecarboxylic acid ester represented by theformula (III-1) is cyclohexanedicarboxylic acid diester and thealicyclic alcohol represented by the formula (III-2) iscyclohexanedimethanol.

According to embodiment III, the alicyclic alcohol is prepared, forexample, by the following process.

The reaction apparatus used is equipped with a preheater for hydrogen, apreheater for the raw materials, a reactor having a shell for heating orcooling by a heat transfer medium, a gas-liquid separator and a hydrogengas circulator. First, a predetermined amount of a preformedcopper-containing catalyst is placed into the reactor, and the catalystis activated in the conventional manner. Then predetermined amounts of acyclohexanecarboxylic acid ester and an aliphatic alcohol having 1 to 4carbon atoms are fed, along with hydrogen gas, to the reactor from itstop while maintaining the reaction system at a specific hydrogenpressure, a specific temperature and a specific superficial linearvelocity of hydrogen gas.

The multitubular reactor described with respect to embodiment II ispreferably used, but according to embodiment III wherein theabove-mentioned specific alcohol is used, a conventional reactor of theshort single-tube type having a shorter tube length and larger tubediameter may also be used.

Examples of the cyclohexanecarboxylic acid ester of the formula (III-1)useful as the raw material include an aromatic ring hydrogenationproduct of an ester prepared by esterifying benzoic acid, terephthalicacid, isophthalic acid, phthalic acid, trimellitic acid, trimesic acid,pyromellitic acid or the like with an aliphatic alcohol having 1 to 4carbon atoms in the conventional manner.

When plural carboxyl groups are attached to the cyclohexane ring ofcyclohexanecarboxylic acid ester, steric configuration thereof may be R-or S-isomer or a mixture of R- and S-isomers.

Specific examples of cyclohexanecarboxylic acid esters of the formula(III-1) are products obtained by hydrogenating the aromatic ring ofesters such as methyl benzoate, ethyl benzoate, n-propyl benzoate,iso-propyl benzoate, n-butyl benzoate, iso-butyl benzoate, dimethylterephthalate, diethyl terephthalate, di-n-propyl terephthalate,di-iso-propyl terephthalate, di-n-butyl terephthalate, di-iso-butylterephthalate, dimethyl isophthalate, diethyl isophthalate, di-n-propylisophthalate, di-iso-propyl isophthalate, di-n-butyl isophthalate,di-iso-butyl isophthalate, dimethyl phthalate, diethyl phthalate,di-n-propyl phthalate, di-iso-propyl phthalate, di-n-butyl phthalate,di-iso-butyl phthalate, trimethyl trimellitate, triethyl trimellitate,tri-n-propyl trimellitate, tri-iso-propyl trimellitate, tri-n-butyltrimellitate, tri-iso-butyl trimellitate, trimethyl trimesate, triethyltrimesate, tri-n-propyl trimesate, tri-iso-propyl trimesate, tri-n-butyltrimesate, tri-iso-butyl trimesate, tetramethyl pyromellitate,tetraethyl pyromellitate, tetra-n-propyl pyromellitate, tetra-iso-propylpyromellitate, tetra-n-butyl pyromellitate, tetra-iso-butylpyromellitate, etc.

Further it is possible to use cyclohexanecarboxylic acid esters havingmixed ester groups such as methyl ethyl terephthalate, methyl butylphthalate, ethyl butyl isophthalate, etc.

Among said cyclohexanecarboxylic acid esters, preferred are commerciallyavailable methyl esters prepared from methanol, and the most preferredare dimethyl 1,4-cyclohexanedicarboxylate and dimethyl1,3-cyclohexanedicarboxylate which are aromatic ring hydrogenationproducts of dimethyl terephthalate and dimethyl isophthalate.

Methods of hydrogenating the aromatic ring are not specifically limitedand include conventional methods such as a method comprisinghydrogenating aromatic carboxylic acid ester in the presence of apreformed supported ruthenium catalyst at a hydrogen pressure of 5 to100 kgf/cm² and at a reaction temperature of about 80 to about 200° C.

The cyclohexanecarboxylic acid ester of the formula (III-1) can also beprepared by esterifying in the conventional manner cyclohexanecarboxylicacid prepared separately with an aliphatic alcohol having 1 to 4 carbonatoms.

Usable as the catalyst for hydrogenation of ester group or groups areconventional preformed copper-containing catalysts capable of reducingesters. Specific examples are copper-chromite, copper oxide/zinc oxide,copper oxide/iron oxide, and these catalysts which contain, as apromoter, oxides of barium, manganese, aluminum, zinc or magnesium.Useful catalysts further include preformed catalysts molded afteraddition of various binders to maintain the strength of catalyst, andsupported catalysts prepared by depositing said oxide on a support suchas alumina, silica, silica-alumina, silicon carbide, zirconia, titaniaor zinc oxide.

Among said copper-containing catalysts, preferred are commerciallyavailable copper-chromite catalysts of the so-called Adkins type,preformed copper oxide/zinc oxide catalysts, and these catalysts whichcontain barium oxide, manganese oxide or the like as a promoter.

Generally, the above copper-chromite catalysts contain copper in anamount, calculated as CuO, of 20-80 wt. %, preferably 30-70 wt. % andchromium in an amount, calculated as Cr₂O₃, of 15-70 wt. %, preferably40-60 wt. %. The above Adkins type catalyst preferably contains copperin an amount, calculated as CuO, of 30-60 wt. % and chromium in anamount, calculated as Cr₂O₃, of 30-60 wt. %. Other copper/metal oxidecatalysts mentioned above preferably contains copper in an amount,calculated as CuO, of 20-95 wt. % and other metal in an amount,calculated as the metal oxide, of 5-80 wt. %. These copper-containingcatalyst preferably contain one or more promoters mentioned above in anamount of up to 10 wt. %, calculated as said metal oxide. In the case ofthe supported catalysts, the percentage values are those calculatedafter excluding the amount of the support.

The shape of the preformed copper-containing catalyst is notspecifically limited, and usually preferred are catalysts of cylindricalshape that are commercially readily available. The size of the preformedcatalyst can be determined according to the interior diameter of thereactor to be used. Usually preferred are cylindrical catalysts having adiameter of 2 to 6 mm and a height of 2 to 6 mm.

In order to control exothermic heat abruptly generated at the start ofreaction or to effectively exhibit a catalyst activity, it is effectivethat the preformed copper-containing catalyst may be subjected topreliminary reduction treatment.

The preliminary reduction treatment can be conducted in the conventionalmanner under a stream of hydrogen-nitrogen gas mixture at an atmosphericpressure or elevated pressure and a temperature of 150 to 300° C. whilegradually increasing the concentration of hydrogen.

Recommended hydrogenation temperature is 200 to 280° C., particularly200 to 270° C., more preferably 220 to 250° C. Below 200° C., the esterused as the raw material is hydrogenated to an insufficient degree,whereas above 280° C., hydrogenolysis occurs, so that the temperature ineither case is undesirable.

A recommendable hydrogen pressure is 185 to 300 kgf/cm², preferably 200to 250 kgf/cm². At a hydrogen pressure of below this range, variousdisadvantages are entailed and include not only reduced reaction rateand decreased productivity, but also increased quantities ofhigh-boiling-point by-products such as ether compounds and wax esters,and lowered yields and selectivity. A lower hydrogen pressure isundesirable also from the viewpoint of duration of activity and strengthof the catalyst. Conversely, a hydrogen pressure of above this range,only the equipment cost increases without appreciably enhancing thereaction rate and selectivity, hence undesirable.

Recommended feed rate of hydrogen is 1 to 40 cm/s, particularly 5 to 30cm/s, and particularly preferable feed rate is 10 to 20 cm/s, in termsof superficial linear velocity. The superficial linear velocity ofhydrogen gas has an extremely marked effect. At less than 1 cm/s,reaction rate is significantly low, making it difficult to achieve thehigh productivity as contemplated in the present invention. A velocityof above 40 cm/s does not provide a further advantage in terms of thereaction, only resulting in increased costs for equipment such as ahydrogen circulator and for energy, hence undesirable.

A preferred feed rate of the ester to be used as the raw material is 1.1to 3.0/h in terms of F/V based on cyclohexanecarboxylic acid ester. Ifthe F/V value is below 1.1/h, productivity is decreased, giving aneconomical disadvantage and the desired alcohol is produced in a loweryield due to side reactions, whereas if the feed rate is above 3.0 /h,the obtained reaction product contains unreacted cyclohexanecarboxylicacid ester, and decreased yield and purity result. Herein, F stands forthe feed rate (liter/h) for feeding the cyclohexanecarboxylic acid esterper hour, and V represents the volume (liter) of catalyst bed in thereactor.

Useful lower aliphatic alcohols to be fed to the reaction systeminclude, for example, C₁₋₄ straight-chain or C₃₋₄ branched-chainaliphatic alcohols such as methanol, ethanol, n-propanol, iso-propanol,n-butanol, iso-butanol, etc. Considering the procedure during therecovery treatment, it is favorable that these alcohols be the same asthe alcohol component (R¹OH) constituting the cyclohexanecarboxylic acidester of the formula (III-1) used as the raw material.

Generally, the hydrogenation of cyclohexane-carboxylic acid ester can berepresented by the following formula (III-3)

wherein R² is a straight- or branched-chain alkyl group having 1 to 4carbon atoms.

However, as the hydrogenation reaction proceeds and the productalicyclic alcohol becomes in excess, an ester interchange reactionoccurs between the cyclohexanecarboxylic acid ester and the producedalicyclic alcohol to achieve an equilibrium relation of the formula(III-4). It is considered that generally a lower alkyl ester ishydrogenated faster than an alicyclic alkyl ester which is shown in theright side of the formula (III-4).

Consequently, if a lower aliphatic alcohol which is the same as theconstituent of the ester used as the raw material is added to thereaction system, the equilibrium relation is considered to move towardthe starting cyclohexanecarboxylic acid lower alkyl ester side, tothereby promote the hydrogenation reaction. The present invention isbased on said mechanism as the background.

The amount of aliphatic alcohol to be fed is preferably 1 to 100% byweight, more preferably 10 to 80% by weight, based on the ester of theformula (III-1) used as the raw material. If the amount is less than 1%by weight, it is difficult to obtain the contemplated result, whereas ifthe alcohol is fed in an amount in excess of 100% by weight, it isdifficult to achieve further remarkably improved effect.

The reaction vessel for the hydrogenation reaction is not specificallylimited, and even a single column type reactor can be used. Amultitubular reactor can also be used which comprises a plurality oftubes having a small interior diameter and arranged in parallel.

As the mode of hydrogenation reaction, there may be mentioned a modecomprising feeding hydrogen and raw material ester and aliphatic alcoholto the top or to the bottom of a fixed-bed reactor. However, whenhydrogen, a raw material and an alcohol are fed from the bottom of thereactor, there arise problems such as decrease in duration of activityand strength of the catallyst due to the use of the lower aliphaticalcohol. Accordingly hydrogen, an ester and aliphatic alcohol as the rawmaterials are preferably fed to the top of the reactor to undergo areaction by the downflow-type process.

When hydrogenation is conducted under said conditions, an alicyclicalcohol can be produced in a high yield. Further, the reaction productcan be purified by conventional methods such as distillation.

Embodiment IV

According to the present inventors' research, it has been found thatwhen a chlorine content of the ruthenium catalyst used for hydrogenatingan aromatic dicarboxylic acid dialkyl ester to producecyclohexane-dicarboxylic acid dialkyl ester is 500 ppm or less, thedesired cyclohexanedicarboxylic acid dialkyl ester is advantageouslyproduced.

Therefore, the desired CHDM is advantageously prepared by using thismethod in the first reaction of the foregoing embodiment I andembodiment II.

Thus, according to the embodiment IV, the cyclohexanedicarboxylic aciddialkyl ester which is an intermediate for preparing CHDM isparticularly prepared by a process for preparing cyclohexanedicarboxylicacid dialkyl ester represented by the formula (IV-2)

wherein R¹ and R² are the same or different and each represents an alkylgroup having 1 to 4 carbon atoms, the process being characterized inthat it comprises the step of hydrogenating an aromatic dicarboxylicacid dialkyl ester represented by the formula (IV-1)

wherein R¹ and R² are as defined above in the presence of a preformedsupported ruthenium catalyst by a fixed-bed continuous reaction, whereinthe preformed supported ruthenium catalyst has a chlorine content of 500ppm or less.

Conventional processes for preparing cyclohexanedicarboxylic aciddialkyl ester include a process comprising carrying out ringhydrogenation of an aromatic dicarboxylic acid dialkyl ester. A typicalexample of such processes comprises ring hydrogenation of a terephthalicacid dialkyl ester in the presence of a preformed supported rutheniumcatalyst by a fixed-bed continuous process to give1,4-cyclohexanedicarboxylic acid dialkyl ester (Japanese UnexaminedPatent Publications Nos. 163554/1979 and 192146/1994).

Ruthenium catalysts are inexpensive compared with palladium catalystsand exhibit a high activity at a low pressure and a low temperature, buthave the drawback of being likely to cause not only the hydrogenation ofaromatic ring but side reactions involving high exothermic heat, such ashydrogenolysis of ester groups to hydroxymethyl group(s) or to methylgroup(s).

Japanese Unexamined Patent Publication No. 163554/1979 discloses atechnique using lithium alumina as a support in an attempt to improvethe duration of the catalyst activity. However, the disclosed techniqueremains to be improved because low productivity is entailed when anaromatic dicarboxylic acid dialkyl ester is used as the raw material.

In the process disclosed in Japanese Unexamined Patent Publication No.192146/1994, hydrogenation is feasible under relatively mild conditions.However, the concentration of terephthalic acid dialkyl ester in thefeed to the reactor is as low as 5 to 20% by weight, and a large amountof reaction product is subjected to the reaction by recycling, resultingin a low yield based on the terephthalic acid dialkyl ester used andleading to a reduced productivity.

In hydrogenating an aromatic dicarboxylic acid dialkyl ester by afixed-bed continuous process using a preformed supported rutheniumcatalyst, this embodiment IV contemplates to provide a novel and usefulprocess for preparing cyclohexanedicarboxylic acid dialkyl ester withsafety at a high productivity on a commercial scale using lesscomplicated production facilities without a likelihood of evolution ofheat owing to side reactions even when the raw material aromaticdicarboxylic acid dialkyl ester is used at a high concentration.

The present inventors conducted extensive research to overcome theforegoing problems and found that when a preformed supported rutheniumcatalyst having specific characteristics is used in ring hydrogenationof said ester, the desired effect can be achieved. Based on thisfinding, the present invention relating to embodiment IV was completed.

The aromatic dicarboxylic acid dialkyl ester to be used in embodiment IVis a dialkyl ester prepared by esterifying terephthalic acid,isophthalic acid or phthalic acid with a monohydric aliphatic alcoholhaving 1 to 4 carbon atoms in the conventional manner.

Specific examples of the monohydric aliphatic alcohol to be used aremethanol, ethanol, n-propanol, iso-propanol, n-butanol, iso-butanol,etc.

Specific examples of the aromatic dicarboxylic acid dialkyl ester to beused are dimethyl terephthalate, diethyl terephthalate, di-iso-propylterephthalate, di-iso-butyl terephthalate, dimethyl isophthalate,diethyl isophthalate, di-iso-propyl isophthalate, di-iso-butylisophthalate, dimethyl phthalate, diethyl phthalate, di-iso-propylphthalate, di-iso-butyl phthalate, etc. Among them, preferred arecommercially readily available dimethyl terephthalate, dimethylisophthalate and/or dimethyl phthalate which are prepared from methanol.

The preformed supported ruthenium catalyst to be used in embodiment IVhas a chlorine content in the preformed supported catalyst in aproportion of up to 500 ppm, preferably 50 to 300 ppm. By adjusting thechlorine content within the above range, reactions other than the ringhydrogenation reaction, namely side reactions, particularlyhydrogenolysis reaction involving cleavage of carbon-oxygen bonds, isadvantageously suppressed so that hydroxymethylcyclohexanecarboxylicacid alkyl ester and methylcyclohexanecarboxylic acid alkyl ester areproduced as by-products in decreased amounts, resulting in a improvedyield of the desired cyclohexanedicarboxylic acid dialkyl ester.

Such supported ruthenium catalysts are readily prepared, for example, bydepositing ruthenium chloride (RuCl₃.3H₂O) on a support in aconventional manner, sufficiently neutralizing the supported catalystwith an alkali and washing the neutralized catalyst with water until thechlorine content becomes up to 500 ppm, followed by drying andreduction.

Among the preformed supported ruthenium catalysts with a chlorinecontent of up to 500 ppm, preferable are those having dispersion ofruthenium, surface distribution thereof and/or pore volume thereof thatfall within specific ranges. That is, preferred catalysts are thosewhich fulfil at least one of the following requirements (i), (ii) and(iii).

(i) having a dispersion of ruthenium of at least 15%, preferably atleast 20%. The preformed supported ruthenium catalyst which is at least15% in the dispersion of ruthenium exhibits a high catalyst activity andachieves high reaction selectivity. The term “dispersion of ruthenium”used herein refers to the percentage of the ruthenium atoms exposed onthe surface of the catalyst, relative to all the atoms of rutheniumdeposited on a support.

(ii) being at least 80 wt. %, preferably at least 90 wt. %, in thesurface distribution of ruthenium. If a catalyst is at least 80 wt. % inthe surface distribution of ruthenium, the catalyst has a higheffectivity of the ruthenium metal on the surface of a support andexhibit an improved catalyst activity, hence more advantageous inpractical use. The term “surface distribution of ruthenium” is usedherein to mean the percentage (weight %) of the ruthenium located nodeeper than 200 μm from the external surface of the support relative tothe total amount of ruthenium.

(iii) being at least 0.20 cc/g, preferably 0.25 to 0.35 cc/g, in thepore volume. If the catalyst has a pore volume of at least 0.20 cc/g,the catalyst is excellent in catalyst activity and in duration ofactivity. Hence such catalyst is improved in practical use. Herein, thepore volume is determined by the mercury intrusion porosimetry.

In short, particularly recommendable preformed supported rutheniumcatalysts are those which are not higher than 500 ppm in chlorinecontent, at least 15% in dispersion of ruthenium, at least 80% by weightin surface distribution of ruthenium and at least 0.20 cc/g in porevolume.

A preferred amount of ruthenium to be deposited on a support is 0.05 to10% by weight, based on the support. If the amount is less than 0.05% byweight, the catalyst activity is low, hence the catalyst is unsuitablefor use. If the amount is more than 10% by weight, it is difficult toimpart a high dispersion of ruthenium, and results in low effectivity ofexpensive ruthenium and markedly low reaction selectivity.

Useful support can be alumina, silica, titania, magnesia, zirconia,silicon carbide and the like, among which alumina is preferred.

As to the shape of the preformed supported ruthenium catalyst,recommended are catalysts of cylindrical shape which are commerciallyreadily available. The size of the preformed catalyst can be suitablyselected according to the interior diameter of the fixed-bed reactor tobe used. Usually preferred are preformed catalysts in the form ofcylinders having a diameter of 2 to 6 mm and a height of 2 to 6 mm.

These preformed supported ruthenium catalysts can be used as such, orcan be used in the reaction after effecting a suitable activationtreatment, such as reduction, in the conventional manner.

As to the recommended mode of the fixed-bed hydrogenation reactionaccording to the present invention, there may be mentioned a gas-liquidmixed phase downflow-type process. The gas-liquid mixed phasedownflow-type process is a process which comprises feeding hydrogen gasand a liquid feed (a melt of raw material or a mixture of a raw materialand a solvent) to the top of a fixed bed reactor packed with theabove-mentioned preformed catalyst, and removing the reaction productand excess hydrogen from the bottom of the reactor, wherein the reactionis carried out under the reaction temperature and hydrogen flow rateconditions such that the reaction temperature is lower than the dewpoint of at least one of the raw material and the product.

Problems are raised if the fixed-bed continuous reaction is conducted byan upflow-type process (wherein the raw material and hydrogen are fedfrom the bottom of the reactor) or by a countercurrent-type process(wherein, for example, the raw material is fed from the top of thereactor and hydrogen is supplied from the bottom of the reactor). Insuch case, the strength of catalyst is adversely affected by thefriction caused by the movement of the catalysts, necessitating frequentreplacement of catalysts. Furthermore, ruthenium metal is separated fromthe preformed supported ruthenium catalyst and is lost, whereby theactivity of the catalyst is reduced within a short period of time.

Also, in the reaction mode wherein the reaction is carried out at areaction temperature which is higher than the dew points of the rawmaterials and the reaction product (gas phase reaction mode),high-boiling-point by-products are deposited on the surface of thecatalyst, significantly reducing the catalyst activity, therebynecessitating frequent replacement of the catalyst and regenerationtreatment thereof. Furthermore, under such conditions, it is necessaryto remove the reaction heat in a gas phase of low heat conductivity sothat complicated equipment is required.

From a commercial viewpoint, it is more preferable to feed to thereactor a raw material (which is normally solid) after conversion to aliquid form by using a solvent, than to feed the raw materials aftermelting it by heating.

Solvents useful for this purpose include, for example, thecyclohexanedicarboxylic acid dialkyl ester produced as the reactionproduct according to the invention, recommendablycyclohexanedicarboxylic acid dialkyl ester corresponding to the aromaticdicarboxylic acid dialkyl ester used as the raw material.

Preferred proportions of aromatic dicarboxylic acid dialkyl ester andcyclohexanedicarboxylic acid dialkyl ester to be used in this case aresuch that the concentration of the former is at least 5% by weight,particularly at least 30% by weight, more preferably at least 40% byweight. At a low concentration of less than 30% by weight, particularlyless than 5% by weight, a low productivity results, and the yield andselectivity tend to decrease due to side reaction ofcyclohexanedicarboxylic acid dialkyl ester circulating through thereaction system.

Among useful solvents other than cyclohexanedicarboxylic acid dialkylester, the alcohol component constituting the aromatic dicarboxylic aciddialkyl ester, i.e. the above-mentioned monohydric aliphatic alcohol, isusable but necessitates the separation and recovery of the solventwithout gaining a particularly good result.

Recommendable feed rate of the aromatic dicarboxylic acid dialkyl esterto be used as the raw material is 0.1 to 5/h, particularly 0.2 to 3/h,and particularly preferable feed rate is 0.5 to 3/h, in terms of F/V(feed rate per hour relative to the volume of the catalyst bed in thereactor). If it is fed at a feed rate of below said range, the reactionachieves lower productivity and is impractical, whereas if it is fed ata feed rate exceeding said range, the resulting reaction productcontains a large amount of unreacted aromatic dicarboxylic acid dialkylester, resulting in significantly low yield and purity, henceundesirable.

The feed rate of hydrogen is preferably 1 to 40 cm/s, particularly 1 to10 cm/s, in terms of superficial linear velocity. In the case of linearvelocity less than 1 cm/s, an effective contact between the gas andliquid can not be achieved on the surface of the catalyst, and thereaction rate decreases and the quantities of by-products of sidereactions increases. On the other hand, at a linear velocity higher than40 cm/s, no further improvement in the reaction is observed, resultingin economical disadvantage, hence undesirable.

Generally, the higher the hydrogen partial pressure is, the moresmoothly the hydrogenation reaction proceeds. Yet, if the hydrogenpressure is higher than necessary, a special pressure-resistant reactoris required, hence economically disadvantageous. Therefore, preferredhydrogen pressure is in the range of 30 to 100 kgf/cm². The use of ahydrogen pressure of less than 30 kgf/cm² results in a low reaction rateand is undesirable. Above 100 kgf/cm², in addition to theabove-mentioned equipment problem, various disadvantages result. Forexample side reactions which involve large amount of exothermic heattend to occur, such as reduction of ester groups or hydrogenolysis ofester groups to methyl group(s). Further, low yields and difficulty inthe control of reaction are entailed.

The reaction crude solution discharged along with excess hydrogen gasfrom the fixed-bed reactor is cooled and separated from the hydrogen gasby a high-pressure gas liquid separator, followed by recovery.

The recovered reaction crude solution is purified by distillation whenso required.

Examples

The following are examples of the embodiments I to IV which illustratethe present invention in further detail. In the following Examples andComparative Examples, the term “L” means “liter”.

First, examples of embodiment I are described.

Example I-1

[First Reaction]

A fixed-bed reactor (20 mm in inner diameter, 1 m in length and 0.314 Lin volume) was charged with 360 g of tableted catalyst (3.2 mm indiameter and 3.2 mm in height) comprising 0.5 wt. % of Ru supported onalumina.

A solution consisting of 30 wt. % of dimethyl terephthalate and 70 wt. %of dimethyl 1,4-cyclohexanedicarboxylate was fed to the top of thereactor at a rate of 628 ml/h (F/V=0.6/h), together with 1.3 Nm³/h ofhydrogen gas (superficial linear velocity under the reactionconditions=4 cm/s), to continuously carry out ring hydrogenation underthe conditions of 140° C. and 40 kg/cm²G.

After carrying out the above fixed-bed continuous ring hydrogenation for10 hours, the the obtained crude reaction product was analyzed by gaschromatography. The composition of the product was as follows.

Dimethyl 1,4-cyclohexanedicarboxylate 96.5 wt. % Low-boiling-pointproduct  2.4 wt. % Methyl 4-hydroxymethylcyclohexane-  0.8 wt. %carboxylate Dimethyl terephthalate  0.2 wt. % High-boiling-point product 0.1 wt. %

[Second Reaction]

A fixed-bed reactor (20 mm in inner diameter, 1 m in length and 0.314 Lin volume) was charged with 490 g of a tableted copper-chromite catalyst(3.5 mm in diameter and 3.5 mm in height) containing barium andmanganese (47 wt. % of copper oxide, 48 wt. % of chromium oxide, 2.5 wt.% of barium oxide and 2.5 wt. % of manganese oxide). The catalyst wasthen subjected to preliminary activation treatment with use of ahydrogen-nitrogen mixed gas.

After the preliminary activation treatment, the crude product of ringhydrogenation in the first reaction was fed to the top of the reactor ata rate of 251 ml/h (F/V=0.8/h) at a temperature of 230° C. and apressure of 200 kg/cm²G, together with 4.9 Nm³/h of hydrogen gas(superficial linear velocity under the reaction conditions=4 cm/s), tocontinuously carry out hydrogenation of ester groups.

The crude reaction product obtained by carrying out hydrogenation of theester groups by the above fixed-bed continuous hydrogenation for 10hours, 1.5 months and 3 months were analyzed for the composition thereofby gas chromatography. The results are shown in Table 1.

Example I-2

The hydrogenation of the ester groups was continuously carried out inthe same manner as in the second reaction in Example I-1 except that thecrude reaction product of the ring hydrogenation was fed at a rate of377 ml/h (F/V=1.2/h) at a reaction temperature of 260° C. The resultsobtained are shown in Table 1.

Comparative Example I-1

The second reaction in Example I-1 was repeated with the exception thathydrogenation of the ring hydrogenation product was carried out at apressure of 150 kg/cm²G. The results of the reaction are shown in Table1.

Comparative Example I-2

The second reaction in Example I-1 was repeated with the exception thatthe hydrogenation of the ring hydrogenation product was carried out at atemperature of 215° C. The results of the reaction are shown in Table 1.

TABLE 1 (wt. %) Example Comp. Example I-1 I-2 I-1 I-2 Low boiling-pointproduct in 10 hours 2.0 2.0 1.8 0.8 in 1.5 months 2.1 2.1 1.7 1.7 in 3months 2.1 2.1 — — HDMT 1) in 10 hours 0.9 0.2 1.4 2.4 in 1.5 months 1.40.4 2.4 3.4 in 3 months 1.5 0.5 — — MOL 2) in 10 hours 1.2 2.2 5.0 5.7in 1.5 months 1.3 2.3 7.2 7.2 in 3 months 1.4 2.4 — — CHDM in 10 hours95.9 95.6 89.6 85.9 in 1.5 months 95.2 95.2 84.2 81.2 in 3 months 95.095.0 — — High boiling-point product in 10 hours trace trace 2.2 5.2 in1.5 months trace trace 4.5 6.5 in 3 months trace trace — — Note: 1)HDMT: Dimethyl 1,4-cyclohexanedicarboxylate 2) MOL: Methyl4-hydroxymethylcyclohexane-carboxylate

According to embodiment I of the present invention, the copper-chromitecatalyst used in the second reaction is remarkably improved in durationof the catalyst activity, and the desired 1,4-cyclohexanedimethanol canbe produced in a high yield with high productivity on a commercialscale.

Hereinafter, the present invention will be described in detail withreference to the following examples of embodiment II.

[Reactor]

The reactor used in the following examples is a multitubularpressure-resistant vessel comprising 15 pressure-resistant tubes U eachhaving an inner diameter of 43 mm and a length of 5 m, and a shell V forheating or cooling the tubes by means of a heat transfer medium isdivided into three zones, wherein the temperature each of the zones areindependently controllable. The raw material and hydrogen fed to the topof the reactor are evenly distributed among the tubes via a distributorW provided in a chamber at the top of the reactor. While monitoring thetemperature with multi-point thermometers set in the tubes, the reactiontemperature was maintained within the desired range by controlling thetemperature of the heat transfer medium in each of the zones. Thecatalyst was uniformly packed into the tubes in such a manner that thetotal volume of the catalyst bed was 100 L. FIG. 1 shows a schematicview of the reactor.

The symbols in FIG. 1 designate the following.

A Raw material tank

B Preheater

C Multitubular reactor

D Heat exchanger

E Gas-liquid separator

F Crude reaction product tank

G Hydrogen gas condenser

H Mist separator

I Hydrogen gas circulator

J-L Heat transfer medium heating/cooling units

M Raw material feed pump

N-P Heat transfer medium circulation pumps

Q-S Cooling medium control valves

T Multi-point thermometer

U Tube

V Shell (for heat transfer medium)

W Distributor

a-s Conduits

Referring to FIG. 1, the reaction will be schematically described.

The raw material is pumped from a raw material tank A to a preheater B,together with hydrogen, by means of a raw material feed pump M through aconduit a. The heated raw material and hydrogen are fed to the top ofthe reactor C through a conduit b.

The crude reaction product and excess hydrogen are conveyed from thebottom of the reactor through a conduit c to a heat exchanger D wherethey are subjected to heat exchange with circulating hydrogen, andconveyed to a gas-liquid separator E through a conduit d. The liquidseparated in this separator enters into a crude reaction product tank Fthrough a conduit e. The gas separated in the separator E passes througha conduit f into a condenser G for cooling, and is conveyed through aconduit g to a mist separator H where the condensate is separated.

The condensate enters into the tank F through a conduit h. The hydrogengas from the mist separator H passes through conduits i and j into ahydrogen gas circulator I, and is conveyed to the heat exchanger Dthrough a conduit k. The hydrogen gas heated in the heat exchanger Denters into the preheater B through a conduit l. The hydrogen gas ispressurized and fed through a conduit s to a conduit j.

The heat transfer medium shell is divided into three zones. In the firstzone at the upper part of the reactor, the heat transfer medium flowingout from the upper part of the first zone passes through a conduit m andis pumped to a heat transfer medium heating/cooling unit J by means of aheat transfer medium circulation pump N. The heat transfer medium cooledor heated in the unit J is fed through a conduit n to the lower part ofthe first zone and circulated. Similarly, in the second and third zones,the heat transfer medium drawn from the upper part thereof is pumped bymeans of a heat transfer medium circulation pump O or P to, and heatedor cooled in, a heating/cooling unit K or L, fed to the lower part ofeach zone and circulated.

In the following examples, crude liquid product of ring hydrogenationreaction (product of the first reaction) was prepared in the firstreaction and stocked in an amount specified. Then, in the secondreaction step, said crude liquid product of the ring hydrogenationreaction was placed in raw material tank A, and the catalyst used in thefirst reaction was replaced by the catalyst (copper-chromite catalyst orthe like) to be used in the second reaction in the multitubular reactorC, and the second reaction was carried out.

However, it is commercially advantageous to carry out the first reactionand second reaction in a continuous manner using a reactor for the firstreaction containing the ruthenium catalyst and a reactor for the secondreaction containing the copper-chromite catalyst or the like.

[Composition Analysis]

The compositions of the raw materials and products of the first andsecond reactions in the examples were analyzed by gas chromatography.

Example II-1

[First Reaction]

The reactor was charged with 95 kg of a tableted catalyst (3.2 mm indiameter and 3.2 mm in height) comprising 1.0 wt. % of Ru supported onalumina. To the top of the reactor, a solution consisting of:

Dimethyl terephthalate 50.0 wt. % Dimethyl 1,4-cyclohexanedicarboxylate48.9 wt. % Methyl 4-hydroxymethylcyclohexane  0.3 wt. % carboxylateMethyl 4-methylcyclohexanecarboxylate  0.8 wt. %

was fed at a rate of 300 L/h(F/V=1.5/h), together with 207 Nm³/h ofhydrogen gas (superficial linear velocity under the reaction conditions=5 cm/s), to continuously carry out ring hydrogenation at a pressure of80 kgf/cm². The reaction temperature was adjusted to 144 to 150° C. inthe upper part of the reactor, 143 to 147° C. in the middle art of thereactor and 139 to 144° C. in the lower part of the reactor. The maximumtemperature difference in the reactor was 11° C.

The composition of the crude liquid reaction product obtained by theabove fixed-bed continuous ring hydrogenation reaction during the periodbetween 5 hours and 15 hours after the start of the reaction is shownbelow.

Dimethyl 1,4-cyclohexanedicarboxylate 97.9 wt. % Methyl4-hydroxymethylcyclohexane  0.6 wt. % carboxylate Methyl4-methylcyclohexanecarboxylate  1.4 wt. % Dimethyl terephthalate  0.1wt. %

[Second Reaction]

The reactor was packed with 155 kg of tableted copper-chromite catalyst(3.5 mm in diameter and 3.5 mm in length) containing barium andmanganese (47 wt. % of copper oxide, 48 wt. % of chromium oxide, 2.5 wt.% of barium oxide and 2.5 wt. % of manganese oxide). The catalyst wasthen subjected to preliminary activation treatment in a stream ofhydrogen-nitrogen mixed gas at atmospheric pressure at a temperature of180 to 200° C., while gradually increasing the hydrogen concentration.

After the preliminary activation treatment, the crude liquid reactionproduct of ring hydrogenation obtained in the first reaction was fed tothe top of the reactor at a feed rate of 188 L/h(F/V=1.84/h) at apressure of 250 kgf/cm², together with 2135 Nm³/h of hydrogen gas(superficial linear velocity under the reaction conditions=20 cm/s), tocontinuously carry out hydrogenation of the ester groups. The reactiontemperature was independently adjusted to 232 to 235° C. in the upperpart of the reactor, 230 to 236° C. in the middle part of the reactorand 228 to 232° C. in the lower part of the reactor.

The reaction product obtained by the above fixed-bed continuoushydrogenation of the ester groups during the period between 5 hours and10 hours after the start of the reaction was of the followingcomposition.

1,4-Cyclohexanedimethanol 97.9 wt. % Dimethyl1,4-cyclohexanedicarboxylate  0.0 wt. % Methyl4-hydroxymethylcyclohexane-  0.0 wt. % carboxylate Low-boiling-pointproduct  2.1 wt. % High-boiling-point product Trace

Example II-2

[First Reaction]

The reactor was charged with 93 kg of a tableted catalyst (3.2 mm indiameter and 3.2 mm in height) comprising 0.5 wt. % of Ru supported onalumina. Dimethyl terephthalate was fed to the top of the reactor at arate of 80 L/h(F/V=0.8/h), together with 170 Nm³/h of hydrogen gas(superficial linear velocity under the reaction conditions=8 cm/s), tocontinuously carry out ring hydrogenation at a pressure of 40 kgf/cm².The reaction temperature was independently adjusted to 146 to 152° C. inthe upper part of the reactor, 138 to 142° C. in the middle part of thereactor and 128 to 132° C. in the lower part of the reactor. The maximumtemperature difference in the reactor was 20° C.

The crude liquid reaction product obtained by the above fixed-bedcontinuous ring hydrogenation reaction during the period between 5 hoursand 10 hours after the start of the reaction was of the followingcomposition.

Dimethyl 1,4-cyclohexanedicarboxylate 97.6 wt. % Methyl4-hydroxymethylcyclohexane  0.5 wt. % carboxylate Methyl4-methylcyclohexanecarboxylate  1.9 wt. % Dimethyl terephthalate  0.0wt. %

[Second Reaction]

Following the procedure of the second reaction in Example II-1, thecrude liquid product of the above first reaction was fed to the top ofthe reactor at a rate of 120 L/h (F/V=1.17/h) at a pressure of 200kgf/cm², together with 855 Nm³/h of hydrogen gas (superficial linearvelocity under the reaction conditions=10 cm/s), to continuously carryout hydrogenation of ester groups. The reaction temperature was adjustedto 232 to 235° C. in the upper part of the reactor, 230 to 236° C. inthe middle part of the reactor and 228 to 232° C. in the lower part ofthe reactor.

The crude reaction product obtained by the above fixed-bed continuoushydrogenation of the ester groups during the period between 5 hours and10 hours after the start of the reaction was of the followingcomposition.

1,4-Cyclohexanedimethanol 97.8 wt. % Dimethyl1,4-cyclohexanedicarboxylate  0.0 wt. % Methyl4-hydroxymethylcyclohexane-  0.0 wt. % carboxylate Low-boiling-pointproduct  2.2 wt. % High-boiling-point product Trace

Example II-3

[First Reaction]

Following the general procedure of the first reaction of Example II-1, asolution consisting of:

Dimethyl terephthalate 50.0 wt. % Dimethyl 1,4-cyclohexanedicarboxylate48.8 wt. % Methyl 4-hydroxymethylcyclo-  0.3 wt. % hexanecarboxylateMethyl 4-methylcyclohexanecarboxylate  0.9 wt. %

was fed to the top of the reactor at a rate of 200 L/h(F/V=1.0/h),together with 192 Nm³/h of hydrogen gas (superficial linear velocityunder the reaction conditions=6 cm/s), to continuously carry out ringhydrogenation at a pressure of 60 kgf/cm². The reaction temperature wasadjusted to 133 to 139° C. in the upper part of the reactor, 133 to 137°C. in the middle part of the reactor and 131 to 135° C. in the lowerpart of the reactor. The maximum temperature difference in the reactorwas 8° C.

The crude liquid reaction product obtained by the above fixed-bedcontinuous ring hydrogenation during the period between 5 hours and 15hours after the start of the reaction was of the following composition.

Dimethyl 1,4-cyclohexanedicarboxylate 98.2 wt. % Methyl4-hydroxymethylcyclohexane-  0.4 wt. % carboxylate Methyl4-methylcyclohexanecarboxylate  1.4 wt. % Dimethyl terephthalate  0.0wt. %

[Second Reaction]

Following the general procedure of the second reaction in Example II-1,the crude product of the above first reaction was fed to the top of thereactor at a rate of 250 L/h(F/V=2.46/h) at a pressure of 250 kgf/cm²,together with 2567 Nm³/h of hydrogen gas (superficial linear velocityunder the reaction conditions=25 cm/s) to continuously carry out thehydrogenation of the ester groups. The reaction temperature was adjustedto 248 to 261° C. in the upper part of the reactor, 246 to 252° C. inthe middle part of the reactor and 248 to 252° C. in the lower part ofthe reactor.

The reaction product obtained by the above fixed-bed continuoushydrogenation of the ester groups during the period between 5 hours and10 hours after the start of the reaction was of the followingcomposition.

1,4-Cyclohexanedimethanol 97.8 wt. % Dimethyl1,4-cyclohexanedicarboxylate  0.1 wt. % Methyl4-hydroxymethylcyclohexane-  0.3 wt. % carboxylate Low-boiling-pointproduct  1.8 wt. % High-boiling-point product Trace

Example II-4

[Second Reaction]

The second reaction in Example II-1 was repeated with the exception thatthe crude product of the following composition obtained by continuingthe first reaction of Example II-1 was used as the raw material and thata copper-chromite catalyst (51 wt. % of copper oxide and 49 wt. % ofchromium oxide) was used. The catalyst was in the form of cylinders witha diameter of 3.5 mm and a height of 3.5 mm.

Dimethyl 1,4-cyclohexanedicarboxylate 97.8 wt. % Methyl4-hydroxymethylcyclohexane-  0.7 wt. % carboxylate Methyl4-methylcyclohexanecarboxylate  1.4 wt. % Dimethyl terephthalate  0.1wt. %

The crude product obtained by the above reaction during the periodbetween 5 hours and 10 hours after the start of the reaction was of thefollowing composition.

1,4-Cyclohexanedimethanol 97.6 wt. % Dimethyl1,4-cyclohexanedicarboxylate  0.0 wt. % Methyl4-hydroxymethylcyclohexane-  0.2 wt. % carboxylate Low-boiling-pointproduct  2.2 wt. % High-boiling-point product Trace

Example II-5

[Second Reaction]

The second reaction in Example II-4 was repeated with the exception thata copper oxide/zinc oxide catalyst (41 wt. % of copper oxide, 50 wt. %of zinc oxide and 9 wt. % of aluminum oxide) was used. The catalyst wasin the form of cylinders with a diameter of 3.5 mm and a height of 3.5mm.

The crude product obtained by the above reaction during the periodbetween 5 hours and 10 hours after the start of the reaction was of thefollowing composition.

1,4-Cyclohexanedimethanol 96.7 wt. % Dimethyl1,4-cyclohexanedicarboxylate  0.4 wt. % Methyl4-hydroxymethylcyclohexane-  0.9 wt. % carboxylate Low-boiling-pointproduct  2.0 wt. % High-boiling-point product Trace

Example II-6

[Second Reaction]

The second reaction in Example II-4 was repeated with the exception thata copper oxide/iron oxide catalyst (30 wt. % of copper oxide, 30 wt. %of iron oxide (FeO) and 40 wt. % of aluminum oxide as support) was used.The catalyst was in the form of cylinders with a diameter of 3.5 mm anda height of 3.5 mm.

The crude product obtained by the above reaction during the periodbetween 5 hours and 10 hours after the start of the reaction was of thefollowing composition.

1,4-Cyclohexanedimethanol 94.2 wt. % Dimethyl1,4-cyclohexanedicarboxylate  1.0 wt. % Methyl4-hydroxymethylcyclohexane-  2.2 wt. % carboxylate Low-boiling-pointproduct  2.5 wt. % High-boiling-point product  0.1 wt. %

Example II-7

[Second Reaction]

The second reaction in Example II-4 was repeated with the exception thata copper oxide/aluminum oxide catalyst (60 wt. % of copper oxide, 6 wt.% of magnesium oxide and 34 wt. % of aluminum oxide as support) wasused. The catalyst was in the form of cylinders with a diameter of 3.5mm and a height of 3.5 mm.

The crude product obtained by the above reaction during the periodbetween 5 hours and 10 hours after the start of the reaction was of thefollowing composition.

1,4-Cyclohexanedimethanol 93.6 wt. % Dimethyl1,4-cyclohexanedicarboxylate  1.4 wt. % Methyl4-hydroxymethylcyclohexane-  3.0 wt. % carboxylate Low-boiling-pointproduct  1.9 wt. % High-boiling-point product  0.1 wt. %

Example II-8

The influence of the length of the tubes in the reactor was determinedin the following manner. The reaction was carried out under the sameconditions as in Example II-1 with the exception that the feed rate ofthe raw material solution was doubled. Then, the obtained reactionproduct was fed again to the reactor under the same conditions. Thus,the reaction corresponding to the tube length of 10 m was carried out.

[First Reaction]

The first reaction in Example II-1 was repeated with the exception thata solution consisting of:

Dimethyl terephthalate 50.0 wt. % Dimethyl 1,4-cyclohexanedicarboxylate48.5 wt. % Methyl 4-hydroxymethylcyclohexane-  0.5 wt. % carboxylateMethyl 4-methylcyclohexanecarboxylate  1.0 wt. %

was fed at a rate of 600 L/h(F/V=3.0/h), and the reaction was carriedout for 5 hours. The obtained product was fed again at a rate of 600 L/hunder the same conditions as in Example II-1 (F/V in total=1.5/h) tocarry out the reaction.

The reaction product obtained by the above reaction during the periodbetween 2 hours and 3 hours after the start of the reaction was of thefollowing composition.

Dimethyl 1,4-cyclohexanedicarboxylate 97.5 wt. % Methyl4-hydroxymethylcyclohexane-  0.4 wt. % carboxylate Methyl4-methylcyclohexanecarboxylate  2.0 wt. % Dimethyl terephthalate  0.1wt. %

[Second Reaction]

The second reaction in Example II-1 was repeated with the exception thatthe above reaction product was used as the raw material and fed at twicethe feed rate of Example II-1, i.e., at 378 L/h(F/V=3.68/h). Thereaction was carried out for 5 hours. The product obtained by thisreaction was fed again at a rate of 378 L/h(F/V in total=1.84/h) underthe same conditions as in the second reaction in Example II-1.

The reaction product obtained by this reaction during the period between2 hours and 3 hours after the start of the reaction was of the followingcomposition.

1,4-Cyclohexanedimethanol 97.6 wt. %  Dimethyl1,4-cyclohexanedicarboxylate 0.0 wt. % Methyl4-hydroxymethylcyclohexane- 0.1 wt. % carboxylate Low-boiling-pointproduct 2.2 wt. % High-boiling-point product Trace

Example II-9

[First Reaction]

Ring hydrogenation was continuously carried out in the same manner as inExample II-2 with the exception that dimethyl isophthalate was used asthe raw material in place of dimethyl terephthalate.

The obtained product was of the following composition.

Dimethyl 1,3-cyclohexanedicarboxylate 97.7 wt. % Methyl3-hydroxymethylcyclohexane- 0.8 wt. % carboxylate Methyl3-methylcyclohexanecarboxylate 1.5 wt. % Dimethyl isophthalate 0.0 wt. %

[Second Reaction]

Ester reduction reaction was continuously carried out in the same manneras in Example II-2 with the exception that the above reaction productwas used as the raw material.

The obtained reaction product was of the following composition.

1,3-Cyclohexanedimethanol 97.3 wt. %  Dimethyl1,3-cyclohexanedicarboxylate 0.0 wt. % Methyl3-hydroxymethylcyclohexane- 0.1 wt. % carboxylate Low-boiling-pointproduct 2.6 wt. % High-boiling-point product Trace

Example II-10

[First Reaction]

The first reaction in Example II-1 was repeated with the exception thata composition of the following make-up was used as the raw material andthat the temperature difference in the tubes in the reactor was changedas described below.

Dimethyl terephthalate 50.0 wt. % Dimethyl 1,4-cyclohexanedicarboxylate48.9 wt. % Methyl 4-hydroxymethylcyclohexane- 0.4 wt. % carboxylateMethyl 4-methylcyclohexanecarboxylate 0.7 wt. %

The temperature of the reactor was adjusted to 175 to 180° C. in theupper part, 136 to 139° C. in the middle part and 118 to 125° C. in thelower part. The maximum temperature difference was 60° C.

The crude liquid product obtained by this fixed-bed continuous reactionduring the period between 5 hours and 10 hours after the start of thereaction was of the following composition.

Dimethyl 1,4-cyclohexanedicarboxylate 86.0 wt. % Methyl4-hydroxymethylcyclohexane- 1.2 wt. % carboxylate Methyl4-methylcyclohexanecarboxylate 7.2 wt. % Dimethyl terephthalate 5.6 wt.%

[Second Reaction]

The second reaction in Example II-1 was repeated with the exception thatthe crude liquid reaction product obtained above was used as the rawmaterial. The reaction product obtained by this fixed-bed continuousester group hydrogenation reaction during the period between 5 hours and10 hours after the start of the reaction was of the followingcomposition.

1,4-Cyclohexanedimethanol 84.7 wt. %

1,4-Cyclohexanedimethanol 84.7 wt. % Dimethyl1,4-cyclohexanedicarboxylate  0.0 wt. % Methyl4-hydroxymethylcyclohexane-  0.1 wt. % carboxylate Low-boiling-pointproduct 15.2 wt. % High-boiling-point product Trace

Example II-11

[First Reaction]

The first reaction in Example II-1 was repeated with the exception thatthe concentration of dimethyl terephthalate in the raw material feed wasdecreased to 20 wt. %. Stated specifically, a solution consisting of:

Dimethyl terephthalate 20.0 wt. % Dimethyl 1,4-cyclohexanedicarboxylate78.2 wt. % Methyl 4-hydroxymethylcyclohexane- 0.5 wt. % carboxylateMethyl 4-methylcyclohexanecarboxylate 1.3 wt. %

was fed to the top of the reactor at a rate of 750 L/h(F/V=1.5/h),together with 207 Nm³/h of hydrogen gas (superficial linear velocityunder the reaction conditions=5 cm/s) to continuously carry out ringhydrogenation at a pressure of 80 kgf/cm². The temperature in the tubeswas adjusted to 144 to 146° C. in the upper part of the reactor, 143 to146° C. in the middle part of the reactor and 140 to 145° C. in thelower part of the reactor. The maximum temperature difference in thereactor was 6° C.

The crude liquid product obtained by this fixed-bed continuous ringhydrogenation reaction during the period between 5 hours and 10 hoursafter the start of the reaction was of the following composition.

Dimethyl 1,4-cyclohexanedicarboxylate 88.5 wt. % Methyl4-hydroxymethylcyclohexane- 1.8 wt. % carboxylate Methyl4-methylcyclohexanecarboxylate 2.1 wt. % Dimethyl terephthalate 7.6 wt.%

[Second Reaction]

The second reaction in Example II-1 was repeated with the exception thatthe above crude liquid product was used as the raw material. Thereaction product obtained by this fixed-bed continuous hydrogenation ofthe ester groups during the period between 5 hours and 10 hours afterthe start of the reaction was of the following composition.

1,4-Cyclohexanedimethanol 88.9 wt. % Dimethyl1,4-cyclohexanedicarboxylate  0.0 wt. % Methyl4-hydroxymethylcyclohexane-  0.1 wt. % carboxylate Low-boiling-pointproduct 11.0 wt. % High-boiling-point product Trace

Comparative Example II-2

[Second Reaction]

The second reaction in Example II-2 was repeated with the exception thatthe same raw material as used in Example II-4 (crude reaction productobtained by continuing the first reaction of Example II-1) was employedand that the reaction pressure was changed to 170 kgf/cm².

The crude product obtained by this reaction during the period between 5hours and 10 hours after the start of the reaction was of the followingcomposition.

1,4-Cyclohexanedimethanol 84.1 wt. % Dimethyl 1,4-cyclohexane-dicarboxylate 0.0 wt. % Methyl 4-hydroxymethylcyclohexane- 9.2 wt. %carboxylate Low-boiling-point product 2.2 wt. % High-boiling-pointproduct 4.5 wt. %

Comparative Example II-3

[Second Reaction]

The second reaction in Example II-2 was repeated with the exception thatthe same raw material as used in Example II-4 was employed and that thehydrogen gas superficial linear velocity was changed to 3 cm/s.

The crude product obtained during the period between 5 hours and 10hours after the start of the reaction was of the following composition.

1,4-Cyclohexanedimethanol 87.8 wt. % Dimethyl1,4-cyclohexanedicarboxylate 1.0 wt. % Methyl4-hydroxymethylcyclohexane- 5.8 wt. % carboxylate Low-boiling-pointproduct 2.2 wt. % High-boiling-point product 3.2 wt. %

Comparative Example II-4

[First Reaction]

The first reaction in Example II-1 was repeated with the exception thatthe raw material feed was fed at four times the feed rate in ExampleII-1. The procedure of feeding the obtained product to the reactor underthe same conditions was repeated three times. Thus, the reactioncorresponding to the tube length of 20 m was carried out. Statedspecifically, a solution consisting of:

Dimethyl terephthalate 50.0 wt. % Dimethyl 1,4-cyclohexanedicarboxylate48.9 wt. % Methyl 4-hydroxymethylcyclohexane- 0.3 wt. % carboxylateMethyl 4-methylcyclohexanecarboxylate 0.8 wt. %

was fed at a rate of 1200 L/h(F/V=6.0/h), and the reaction was carriedout for about 5 hours under the same condition as in Example II-1. Theobtained product was fed again to the reactor at a rate of 1200 L/hunder the same conditions as in Example II-1. The procedure of feedingthe obtained product to the reactor under the same conditions wasrepeated.

The reaction product obtained in the fourth reaction of the aboveprocedure (F/V in total=1.5/h) during the period between 2 hours and 3hours after the start of the reaction was of the following composition.

Dimethyl 1,4-cyclohexanedicarboxylate 89.4 wt. % Methyl4-hydroxymethylcyclohexane- 0.8 wt. % carboxylate Methyl4-methylcyclohexanecarboxylate 3.5 wt. % Dimethyl terephthalate 6.3 wt.%

[Second Reaction]

The second reaction was carried out in the same manner as in the firstreaction to effect the reaction corresponding to the tube length of 20m. Thus, the reaction was carried out for 5 hours using the same rawmaterial as used in Example II-4 under the same conditions as in thesecond reaction of Example II-1 with the exception that the raw materialwas fed at four times the feed rate in Example II-1, i.e., at 752 L/h(F/V=7.35/h). The obtained reaction product was fed again to the reactorat a rate of 752 L/h under the same conditions as in the second reactionof Example II-1. The procedure of feeding the obtained product to thereactor under the same conditions was repeated three times.

The reaction product obtained in the fourth reaction of the aboveprocedure (F/V in total=1.84/h) during the period between 2 hours and 3hours after the start of the reaction was of the following composition.

1,4-Cyclohexanedimethanol 84.8 wt. % Dimethyl1,4-cyclohexanedicarboxylate 3.5 wt. % Methyl4-hydroxymethylcyclohexane- 7.8 wt. % carboxylate Low boiling-pointproduct 1.8 wt. % High-boiling-point product 2.1 wt. %

Comparative Example II-1

The reactor used was a pressure-resistant column reactor with an innerdiameter of 159 mm and a length of 2 m, which comprises a shell forheating and cooling by a heat transfer medium. The raw material andhydrogen fed to the top of the reactor pass through perforated plateprovided in the upper part of the reactor and are distributed throughthe zone to be charged with a catalyst.

[First Reaction]

The reactor was packed with 35 L of tableted catalyst (3.2 mm indiameter and 3.2 mm in height) comprising 1.0 wt. % of Ru supported onalumina. To the top of the reactor, a solution consisting of:

Dimethyl terephthalate 50.0 wt. % Dimethyl 1,4-cyclohexanedicarboxylate48.5 wt. % Methyl 4-methylcyclohexanecarboxylate 1.0 wt. % Methyl4-hydroxymethylcyclohexane- 0.5 wt. % carboxylate

was fed at a rate of 105 L/h(F/V=1.5/h), together with 182 Nm³/h ofhydrogen gas (superficial linear velocity under the reactionconditions=5 cm/s), to continuously carry out ring hydrogenation at apressure of 80 kgf/cm². The reaction temperature was measured with athermometer set in the central part of the reactor, and found to be 144to 178° C.

The crude reaction product obtained by the above fixed-bed continuousring hydrogenation during the period between 3 hours and 5 hours afterthe start of the reaction was of the following composition.

Dimethyl 1,4-cyclohexanedicarboxylate 79.9 wt. % Methyl4-hydroxymethylcyclohexane- 3.2 wt. % carboxylate Methyl4-methylcyclohexanecarboxylate 7.5 wt. % Dimethyl terephthalate 2.6 wt.% Low-boiling-point product 6.8 wt. %

[Second Reaction]

The reactor of the same type as in the above first reaction was chargedwith 35 L of tableted copper-chromite catalyst (3.5 mm in diameter and3.5 mm in height) containing barium and manganese (47 wt. % of copperoxide, 48 wt. % of chromium oxide, 2.5 wt. % of barium oxide and 2.5 wt.% of manganese oxide). The catalyst was then subjected to preliminaryactivation treatment under the same conditions as in Example II-1. Afterthe preliminary activation treatment, the raw material of the samecomposition as in Example II-4 was fed to the top of the reactor at arate of 66 L/h(F/V=1.84/h) and a pressure of 250 kgf/cm², together with1916 Nm³/h of hydrogen gas (superficial linear velocity under thereaction conditions=20 cm/s), to continuously carry out hydrogenation ofester groups at 230 to 245° C.

The reaction product obtained by the above reaction during the periodbetween 5 hours and 6 hours after the start of the reaction was of thefollowing composition.

1,4-Cyclohexanedimethanol 75.7 wt. % Dimethyl1,4-cyclohexanedicarboxylate 5.8 wt. % Methyl4-hydroxymethylcyclohexane- 11.5 wt. % carboxylate Low-boiling-pointproduct 2.8 wt. % High-boiling-point product 4.2 wt. %

The embodiment II of the present invention provides a production processwhich is capable of remarkably improving the productivity per reactorand giving high-quality cyclohexanedimethanol by a simplified procedurein a high yield on a commercial scale.

The following are examples of embodiment III.

[Reactor]

The reactor used in the following examples was a pressure-resistantvessel with an internal diameter of 30 mm, a length of 3 m and aninternal volume of 2100 ml. The temperature of the reactor wasadjustable by heating or cooling with a heat transfer medium. Thereaction temperature was measured with a multi-point thermometerprovided in the reactor.

[Composition Analysis]

The compositions of the ester as the raw material and the reactionproducts in the following examples were analyzed by gas chromatography.

Example III-1

The reactor was packed with 2.8 kg of tableted copper-chromite catalyst(3.5 mm in diameter and 3.5 mm in height) containing barium andmanganese (46 wt. % of copper oxide, 49 wt. % of chromium oxide, 2.4 wt.% of barium oxide and 2.6 wt. % of manganese oxide). The catalyst wasthen subjected to a preliminary activation treatment in a stream ofhydrogen-nitrogen mixed gas at atmospheric pressure and at a temperatureof 150 to 200° C. while gradually increasing the hydrogen concentration.

After the preliminary activation treatment, 2.7 kg/h of dimethyl1,4-cyclohexanedicarboxylate with a purity of 99.5% (containing 0.5% ofmethyl 4-hydroxymethylcyclohexanecarboxylate) and 270 g/h of methanol(10 wt. % relative to the starting ester) were fed to the top of thereactor at a hydrogen pressure of 200 kgf/cm², together with 27 Nm³/h ofhydrogen gas (superficial linear velocity under the reactionconditions=10 cm/s). Thus, hydrogenation of the ester groups wascontinuously carried out.

The reaction temperature was adjusted to 236 to 240° C. in the upperpart of the reactor, 239 to 241° C. in the middle part of the reactorand 237 to 238° C. in the lower part of the reactor. The reactionproduct obtained by this fixed-bed continuous hydrogenation reactionduring the period between 5 hours and 10 hours after the start of thereaction was of the following composition.

1,4-Cyclohexanedimethanol 95.2 wt. %  Dimethyl1,4-cyclohexanedicarboxylate Trace Methyl 4-hydroxymethylcyclohexane-0.3 wt. % carboxylate Low-boiling-point product 2.1 wt. %High-boiling-point product 2.4 wt. %

Example III-2

Hydrogenation of ester groups was continuously carried out in the samemanner as in Example III-1 with the exception that methanol was fed at arate of 810 g/h(30 wt. % relative to the starting ester). The reactionproduct obtained by the fixed-bed continuous hydrogenation during theperiod between 5 hours and 10 hours after the start of the reaction wasof the following composition.

1,4-Cyclohexanedimethanol 97.4 wt. % Dimethyl1,4-cyclohexanedicarboxylate 0.0 wt. % Methyl4-hydroxydimethylcyclohexane- 0.1 wt. % carboxylate Low-boiling-pointproduct 1.5 wt. % High-boiling-point product 1.0 wt. %

Example III-3

Hydrogenation of ester groups was continuously carried out in the samemanner as in Example III-2 with the exception that the reactiontemperature was adjusted to 228 to 230° C. in the upper part of thereactor, 229 to 231° C. in the middle part of the reactor and 227 to229° C. in the lower part of the reactor. The reaction product obtainedby the fixed-bed continuous hydrogenation reaction during the periodbetween 5 hours and 10 hours after the start of the reaction was of thefollowing composition.

1,4-Cyclohexanedimethanol 97.5 wt. % Dimethyl1,4-cyclohexanedicarboxylate 0.0 wt. % Methyl4-hydroxymethylcyclohexane- 0.4 wt. % carboxylate Low-boiling-pointproduct 0.9 wt. % High-boiling-point product 1.2 wt. %

Example III-4

Hydrogenation of ester groups was continuously carried out in the samemanner as in Example III-1 with the exception that methanol was fed at arate of 2160 g/h(80 wt. % relative to the starting ester). The reactionproduct obtained by the fixed-bed continuous hydrogenation reactionduring the period between 5 hours and 10 hours after the start of thereaction was of the following composition.

1,4-Cyclohexanedimethanol 98.9 wt. % Dimethyl1,4-cyclohexanedicarboxylate 0.0 wt. % Methyl4-hydroxymethylcyclohexane- 0.1 wt. % carboxylate Low-boiling-pointproduct 0.7 wt. % High-boiling-point product 0.3 wt. %

Example III-5

Hydrogenation of ester groups was continuously carried out in the samemanner as in Example III-1 with the exception that methanol was fed at arate of 2700 g/h (100 wt. % relative to the starting ester). Thereaction product obtained by the fixed-bed continuous hydrogenationreaction during the period between 5 hours and 10 hours after the startof the reaction was of the following composition.

1,4-Cyclohexanedimethanol 98.7 wt. % Dimethyl1,4-cyclohexanedicarboxylate 0.0 wt. % Methyl4-hydroxymethylcyclohexane- 0.2 wt. % carboxylate Low-boiling-pointproduct 0.7 wt. % High-boiling-point product 0.4 wt. %

Example III-6

Hydrogenation of ester groups was continuously carried out in the samemanner as in Example III-1 with the exception that methanol was notused. The reaction product obtained by the fixed-bed continuoushydrogenation reaction during the period between 5 hours and 10 hoursafter the start of the reaction was of the following composition.

1,4-Cyclohexanedimethanol 86.2 wt. % Dimethyl1,4-cyclohexanedicarboxylate 0.5 wt. % Methyl4-hydroxymethylcyclohexane- 2.2 wt. % carboxylate Low-boiling-pointproduct 5.1 wt. % High-boiling-point product 6.0 wt. %

In this Example III-6, no alcohol was used so that the yield of1,4-cyclohexanedimethanol is lower than those of Examples III-1 toIII-5.

In Example II-1 according to embodiment II, the multitubular reactor(comprising 15 tubes having a diameter of 43 mm and a length of 5 m) sothat the high yield can be achieved even without use of the alcohol.However, when a single-column reactor having a relatively short lengthof 3 m as in the above Example III-6, the residence time within thereactor is short and therefore there is a tendency that the yieldbecomes lowered.

The production process according to embodiment III of the presentinvention is a process wherein low-boiling-point and high-boiling-pointproducts are produced in reduced amounts by using a simple apparatus,namely without using a complicated reactor such as a multitubularreactor. Further, alicyclic alcohols, in particularcyclohexanedimethanol, can be produced in high yields. Therefore, theprocess of the present invention is very advantageous for the commercialpurpose.

The present invention will be described below in detail with referenceto examples of embodiment IV, wherein the characteristics of thecatalysts were tested and evaluated in the following manners.

Chlorine content (%): Measured by titrimetric method using silvernitrate. The end point for the titration was determined by checkingexcess silver nitrate with an indicator (potassium chromate) for thecolor change point (from yellow to faint red).

Dispersion: Determined by the gas adsorption method.

Surface distribution (%): Determined with an X-ray microanalyzer (EPMA).

Pore volume: Determined with a mercury intrusion porosimetry.

The catalysts used in the following examples were tableted catalyst (3.2mm in diameter and 3.2 mm in height) comprising 0.5 wt. % of Rusupported on alumina.

Example IV-1

A fixed-bed continuous reactor (20 mm in inner diameter, 5 m in lengthand 1.57 liters in volume) was used, in which three parts thereof,namely the upper, middle and lower parts thereof, can be independentlyheated or cooled with respective electric heaters or cooling fans. Thereactor was charged with 1.83 kg of tableted catalyst (3.2 mm indiameter and 3.2 mm in height) comprising 0.5 wt. % of Ru supported onalumina, the catalyst having the characteristics shown in Table 2.

To the top of the reactor, a raw material solution consisting of:

Dimethyl terephthalate (hereinafter “DMT”) 50 wt. % Dimethyl1,4-cyclohexanedicarboxylate 48.7 wt. % (hereinafter “HDMT”) Methyl4-hydroxymethylcyclohexane- 0.3 wt. % carboxylate (hereinafter “MOL”)Methyl 4-methylcyclohexanecarboxylate 1.0 wt. % (hereinafter “MME”)

was fed at a rate of 3.7 L/h(F/V=1.18/h), together with 1.46 Nm³/h ofhydrogen gas (superficial linear velocity under the reactionconditions=4 cm/s), to continuously carry out ring hydrogenation at apressure of 50 kgf/cm².

The reaction temperature was 144 to 148° C. in the upper part of thereactor, 143 to 146° C. in the middle part of the reactor and 140 to145° C. in the lower part of the reactor. The fixed-bed continuous ringhydrogenation was carried out for 10 hours, and composition of theobtained crude reaction product was analyzed by gas chromatography. Theresults are shown in Table 2.

Example IV-2

The reactor of the same type as in Example IV-1 was charged with 1.85 kgof preformed Ru-supported catalyst having the characteristics shown inTable 2. The raw material solution described in Example IV-1 was fed tothe top of the reactor at a rate of 2.7 L/h (F/V=0.86/h), together with2.85 Nm³/h of hydrogen gas (superficial linear velocity under thereaction conditions=6 cm/s), to continuously carry out ringhydrogenation at a pressure of 65 kgf/cm².

The reaction temperature was 152 to 154° C. in the upper part of thereactor, 145 to 149 ° C. in the middle part of the reactor and 143 to147° C. in the lower part of the reactor. The fixed-bed continuous ringhydrogenation was carried out for 10 hours, and the composition of theobtained crude reaction product was analyzed by gas chromatography. Theresults are shown in Table 2.

Example IV-3

Ring hydrogenation was carried out in the same manner as in Example IV-2with the exception of using 1.79 kg of tableted supported Ru catalysthaving the characteristics shown in Table 2. The reaction temperaturewas 151 to 155° C. in the upper part of the reactor, 145 to 151° C. inthe middle part of the reactor and 143 to 149° C. in the lower part ofthe reactor. The fixed-bed continuous ring hydrogenation was carried outfor 10 hours, and the composition of the obtained reaction product wasanalyzed by gas chromatography. The results are shown in Table 2.

Example IV-4

Ring hydrogenation was carried out in the same manner as in Example IV-2with the exception of using 1.86 kg of tableted supported Ru catalystwith the characteristics shown in Table 2. The reaction temperature was152 to 155° C. in the upper part of the reactor, 146 to 151° C. in themiddle part of the reactor and 144 to 149° C. in the lower part of thereactor. The fixed-bed continuous ring hydrogenation was carried out for10 hours, and the composition of the obtained crude reaction product wasanalyzed by gas chromatography. The results are shown in Table 2.

Example IV-5

Ring hydrogenation was carried out in the same manner as in Example IV-2with the exception of using 1.88 kg of tableted supported Ru catalystwith the characteristics shown in Table 2. The reaction temperature was152 to 155° C. in the upper part of the reactor, 146 to 150° C. in themiddle part of the reactor and 144 to 149° C. in the lower part of thereactor. The fixed-bed continuous ring hydrogenation was carried out for10 hours, and the composition of the obtained crude reaction product wasanalyzed by gas chromatography. The results are shown in Table 2.

Comparative Example IV-1

Ring hydrogenation was carried out in the same manner as in Example IV-2with the exception of using 1.83 kg of tableted supported Ru catalystwith the characteristics shown in Table 2. The reaction temperature was152 to 153° C. in the upper part of the reactor, 146 to 150° C. in themiddle part of the reactor and 144 to 148° C. in the lower part of thereactor. The above fixed-bed continuous ring hydrogenation was carriedout for 10 hours, and the composition of the obtained crude reactionproduct was analyzed by gas chromatography. The results are shown inTable 2.

Example IV-6

Ring hydrogenation was carried out in the same manner as in Example IV-2with the exception that tableted supported Ru catalyst (chlorine contentof 205 ppm, dispersion of 22%, surface distribution of 90 wt. % and porevolume of 0.33 cc/g) was used in an amount of 1.76 kg and that asolution consisting of:

Dimethyl isophthalate 50.0 wt. % Dimethyl 1,3-cyclohexanedicarboxylate48.2 wt. % Methyl 3-hydroxymethylcyclohexane- 0.5 wt. % carboxylateMethyl 3-methylcyclohexanecarboxylate 1.3 wt. %

was used as the raw material. The reaction temperature was 150 to 153°C. in the upper part of the reactor, 146 to 148° C. in the middle partof the reactor and 144 to 147° C. in the lower part of the reactor. Thefixed-bed continuous ring hydrogenation was carried out for 10 hours,and the composition of the obtained crude reaction product was analyzedby gas chromatography. The results are shown below.

Dimethyl 1,3-cyclohexanedicarboxylate 97.9 wt. % Methyl3-hydroxymethylcyclohexane- 0.7 wt. % carboxylate Methyl3-methylcyclohexanecarboxylate 1.3 wt. % Dimethyl isophthalate 0.1 wt. %

TABLE 2 Comp. Example Ex. IV-1 IV-2 IV-3 IV-4 IV-5 IV-1 Characteristicsof catalyst Chlorine content 230 180 230 220 230 670 (ppm) Dispersion(%) 25 31 11 25 20 25 Surface 89 94 85 71 81 92 distribution (wt %) Porevolume (cc/g) 0.30 0.35 0.31 0.33 0.15 0.32 Reaction product composition(wt. %) HDMT 97.9 98.1 95.5 95.4 94.2 92.3 DMT 0.1 0.2 1.8 2.2 2.2 0.3MOL 0.6 0.4 1.2 1.1 1.4 2.4 MME 1.4 1.3 1.5 1.3 2.2 5.0

Embodiment IV of the present invention provides a simple andcommercially advantageous process for producing cyclohexanedicarboxylicacid dialkyl esters using a special preformed supported Ru catalyst, theprocess being capable of giving high-quality cyclohexane-dicarboxylicacid dialkyl ester in a high yield with remarkably improved productivityand safety.

Therefore, when a cyclohexanedicarboxylic acid dialkyl ester is preparedby such a method, the use of such cyclohexanedicarboxylic acid dialkylester as a raw material can advantageously give the desiredcyclohexanedimethanol.

What is claimed is:
 1. A process for preparing cyclohexanedimethanolrepresented by the formula (1):

the process comprising the steps of (a) ring hydrogenating aromaticdicarboxylic acid dialkyl ester represented by the formula (3)

 (wherein R is an alkyl group having 1 to 4 carbon atoms), in thepresence of a preformed supported ruthenium catalyst by a fixed-bedcontinuous reaction to give a cyclohexanedicarboxylic acid dialkyl esterrepresented by the formula (2):

 (wherein R is as defined above), and (b) hydrogenating thecyclohexanedicarboxylic acid dialkyl ester obtained in the step (a) andrepresented by the formula (2) by a fixed-bed continuous reaction in thepresence of a preformed copper-containing catalyst under the conditionsof reaction temperature of 225 to 280° C., hydrogen pressure of 200 to250 kgf/cm² and hydrogen gas feed rate of 2 to 20 cm/s in terms ofsuperficial linear velocity and the feed rate of thecyclohezanedicarboxylic acid dialkyl ester per hour per unit volume ofthe catalyst bed in the reactor of 0.2 to 2.0/h., wherein both of step(a) and step (b) are of a single column type reactor.
 2. The process forpreparing cyclohexanedimethanol according to claim 1, wherein thecyclohexanedimethanol of the formula (1) is 1,4-cyclohexanedimethanol;the aromatic dicarboxylic acid dialkyl ester of the formula (3) isterephthalic acid dialkyl ester; the cyclohexanedicarboxylic aciddialkyl ester of the formula (2) is 1,4-cyclohexanedicarboxylic aciddialkyl ester; and the preformed copper-containing catalyst is apreformed copper chromite catalyst.
 3. The process for preparingcyclohexanedimethanol according to claim 2, wherein the terephthalicacid dialkyl ester is dimethyl terephthalate.
 4. The process forpreparing cyclohexanedimethanol represented by the formula (1):

the process comprising the steps of (a) continuously feeding thearomatic dicarboxylic acid dialkyl ester and hydrogen to the top of amultitubular pressure-resistant reactor packed with the preformedsupported ruthenium catalyst to effect hydrogenation under a gas-liquidmixed phase condition, and removing excess hydrogen and thecorresponding cyclohexanedicarboxylic acid dialkyl ester from the bottomof said reactor, wherein each tube of the multitubular reactor has aninner diameter of 2.5 to 10 cm and a length of 3 to 15 m, and (b)continuously feeding the cyclohexanedicarboxylic acid dialkyl esterobtained in step (a) above and hydrogen to the top of a multitubularpressure-resistant reactor packed with the preformed copper-containingcatalyst to effect hydrogenation under a gas-liquid mixed phasecondition, wherein reaction temperature is 200 to 280° C., hydrogenpressure is 185 to 300 kgf/cm² and hydrogen gas feed rate is 5 to 30cm/s in terms of superficial linear velocity, and removing excesshydrogen and the resulting cyclohexanedimethanol from the bottom of saidreactor, wherein each tube of the multitubular reactor has an innerdiameter of 2.5 to 10 cm and a length of 3 to 5 m, wherein the number oftubes is 10 to 2000 in both of step (a) and step (b).
 5. The process forpreparing cyclohexanedimethanol according to claim 4, wherein in step(a), the aromatic dicarboxylic acid dialkyl ester is dimethylterephthalate, dimethyl isophthalate or dimethyl phthalate.
 6. Theprocess for preparing cyclohexanedimethnaol according to claim 4,wherein in step (a), the aromatic dicarboxylic acid dialkyl ester is fedas it is, or fed as admixed with a reaction product of step (a)predominantly comprising the cyclohexanedicarboxylic acid dialkyl ester;and wherein the hydrogenation in step (a) is carried out at a reactiontemperature of 120 to 180° C., hydrogen pressure of 30 to 100 kgf/cm²and hydrogen gas feed rate of 1 to 15 cm/s in terms of superficiallinear velocity.
 7. The process for preparing cyclohexanedimethanolaccording to claim 6, wherein the aromatic dicarboxylic acid dialkylester is fed as admixed with the reaction product of step (a), and theconcentration of the aromatic dicarboxylic acid dialkyl ester in themixture is at least 30% by weight.
 8. The process for preparingcyclohexanedimethnaol according to claim 4, wherein in step (b), thecyclohexanedicarboxylic acid dialkyl ester is fed as it is, or fed asadmixed with a reaction product of step (b) predominantly comprising thecyclohexanedimethnaol, the concentration of the cyclohexanedicarboxylicacid dialkyl ester in the mixture being at least 90 weight %, andwherein the hydrogenation in step (b) is carried out under theconditions of reaction temperature of 200 to 280° C., hydrogen pressureof 185 to 300 kgf/cm² and hydrogen gas feed rate of 5 to 30 cm/s interms of superficial linear velocity.
 9. The process for preparingcyclohexanedimethanol according to claim 4, wherein in step (b), thefeed rate of the cyclohexanedicarboxylic acid dialkyl ester per hour perunit volume of the catalyst bed in the multitubular pressure-resistantreactor is 1.1 to 3.0/h.
 10. The process for preparingcyclohexanedimethanol according to claim 4, wherein the hydrogenation instep (a) is carried out while controlling the temperature differencewithin the multitubular reactor to not greater than 50° C.
 11. Theprocess for preparing cyclohexanedimethanol according to claim 4,wherein the multitubular reactor used in step (a) comprises a shell anda plurality of tubes contained in the shell, wherein the shell isdivided into at least two separate zones, through each of which a heattransfer medium is circulated for heating or cooling the tubes, each ofthe heat transfer media being independently heated or cooled in such amanner that the temperature difference within the reactor is controlledto not greater than 50° C.
 12. The process for preparingcyclohexandimethanol represented by the formula (1):

the process comprising the steps of (a) ring hydrogenating aromaticdicarboxylic acid dialkyl ester represented by the formula (3)

 (wherein R is an alkyl group having 1 to 4 carbon atoms) in thepresence of a preformed supported ruthenium catalyst by a fixed-bedcontinuous reaction to give a cyclohexanedicarboxylic acid dialkyl esterrepresented by the formula (2):

 (wherein R is as defined above), and (b) hydrogenating thecyclohexanedicarboxylic acid dialkyl ester obtained in the step (a) andrepresented by the formula (2) by a fixed-bed continuous reaction in thepresence of preformed copper-chromite catalyst under the conditions ofreaction temperature of 200 to 280° C., hydrogen pressure of 185 to 300kgf/cm² and hydrogen gas feed rate of 5 to 30 cm/s in terms ofsuperficial linear velocity, wherein in step (b), hydrogen, thecyclohexanedicarboxylic acid dialkyl ester and an aliphatic alcoholhaving 1 to 4 carbon atoms and an aliphatic alcohol having 1 to 4 carbonatoms are continuously fed to the top of a reactor packed with thepreformed copper-containing catalyst in an amount of 10 to 80% by weightbased on the cyclohexanedicarboxylic acid dialkyl ester to carry outhydrogenation under a gas-liquid mixed phase condition, and excesshydrogen, the reaction product and the aliphatic alcohol having 1 to 4carbon atoms are removed from the bottom of the reactor.